Method for producing acrolein and/or acrylic acid

ABSTRACT

Acrolein and/or acrylic acid are prepared from propane and/or propene by a process comprising the following steps: (a) separation of propane and/or propene from a propane- and/or propene-containing gas mixture by absorption in an absorbent, (b) separation of the propane and/or propene from the absorbent to give a gas B and (c) use of the gas B obtained in stage (b) for an oxidation of propane and/or propene to acrolein and/or acrylic acid, no heterogeneously catalyzed dehydrogenation of propane without supply of oxygen being carried out between steps (b) and (c).

The present invention relates to a process for the preparation ofacrolein and/or acrylic acid from propane and/or propene.

Acrolein and acrylic acid are key chemicals. Thus, acrylic acid is used,inter alia, as a monomer for the preparation of polymers which, forexample as a dispersion in an aqueous medium, are used as binders.Depending on the intended use of the polymer, an esterification of theacrylic acid may also take place before the polymerization. Acrolein isan important intermediate, for example for the preparation ofglutaraldehyde, methionine, folic acid and acrylic acid.

Known processes for the preparation of acrolein and/or acrylic acidstart from propane and/or propene. DE-A 33 13 573 and EP-A-0 117 146disclose a process for converting propane into acrolein and/or acrylicacid. A two-stage or three-stage process in which the propane isdehydrogenated to propene in a first stage and the propene is oxidizedto acrolein in a second stage is described. An important feature here isthat no separation of the propane from secondary components formed inthe dehydrogenation, for example molecular hydrogen, is carried outbetween the two stages. The oxidation is carried out under conditionsunder which no marked oxidation of the hydrogen takes place. In a thirdstage, the acrolein can then be oxidized to acrylic acid. It is alsopossible to separate off unconverted propane and propene from the secondor third stage by absorption and, after liberation from the absorbent,to recycle them to the first stage (dehydrogenation stage).

Japanese Patent Application JP-A 10-36311 discloses a process for thepreparation of α,β-unsaturated carboxylic acids, such as acrylic acid,by gas-phase catalytic oxidation of propane in the presence of acomposite metal oxide catalyst, the ratio of propane to oxygen and, ifrequired, diluent gas being kept in a defined range for achieving highyields in the starting mixture and furthermore the conversion being keptat a specific value. Unconverted propane can be separated off after theoxidation by a selective separator which comprises pressure swingadsorption units and then recycled to the gas-phase catalytic oxidation.

GB 1378-178 discloses a process in which unconverted hydrocarbon from anoxidation process is absorbed in an absorbent and the absorbent isstripped with a stripping medium. In this, the hydrocarbon to berecovered is added to the stripping medium in such a quantity that themixture is outside the flammable limits.

It is an object of the present invention to provide a process for thegas-phase catalytic preparation of acrolein and/or acrylic acid frompropane and/or propene, which is economical and in which the catalystused can be employed for a very long time without regeneration.

We have found that this object is achieved, according to the invention,by absorption of propane and/or propene from a propane- and/orpropene-containing gas mixture into an absorbent, separation of thepropane and/or propene from the absorbent and subsequent use of thepropane and/or propene for an oxidation to acrolein and/or acrylic acid.

The present invention therefore relates to a process for the preparationof acrolein and/or acrylic acid from propane and/or propene, the processcomprising the following steps:

-   -   (a) separation of propane and/or propene from propane- and/or        propene-containing gas mixture A by absorption into an        absorbent,    -   (b) separation of the propane and/or propene from the absorbent        to give a propane- and/or propene-containing gas B and    -   (c) use of the gas B obtained in step (b) for an oxidation of        propane and/or propene to acrolein and/or acrylic acid,        no heterogeneously catalyzed dehydrogenation of propane without        a supply of oxygen being carried out between step (b) and step        (c). Preferred embodiments of the invention are evident from the        following description and the figures.

Since the propane and/or propane are subjected to absorption before theoxidation, as a rule residues of absorbent are present in the gas B.Surprisingly, it has now been found that nevertheless no problems occurduring the oxidation. Thus, no substantial decrease in the activity ofthe oxidation catalyst was observed, and the oxidation catalyst could beused over a long operating period without regeneration. Furthermore, noproblems due to any expected concomitant oxidation of residues ofabsorbent in the oxidation stage were observed. Where residues of theabsorbent present problems, which is generally not the case whenhydrocarbons having a high boiling point, in particular paraffins, areused as the absorbent, said absorbent can be removed, for example byquenching with water or by adsorption.

In the process according to DE-A 33 13 573, propane and propenerecovered by absorption or separated off are recycled to theheterogeneously catalyzed dehydrogenation of propane. In theheterogeneously catalyzed propane dehydrogenation, the catalyst may bedeactivated, for example by coking. Such dehydrogenation catalysts aretherefore frequently regenerated. The absorbent fed in together with thegas stream therefore presents no problems since it can be incineratedtogether with the coke. On the other hand, the catalysts used in theoxidation to acrolein and/or acrylic acid are usually not regenerated sooften because the additional regeneration cost which arises by virtue ofthe fact that the feed gas contains absorbent is greater than in thecase of dehydrogenation. The novel process has the advantage that theoxidation catalyst can be used over a long period without regeneration.

The novel process differs from the process according to DE-A 33 13 573in that the propane and/or propene separated off by absorption is/arefed to an oxidation stage. In contrast to the situation in the processaccording to DE-A 33 13 573, there is, according to the invention, noheterogeneously catalyzed dehydrogenation of propane without supply ofoxygen between the separation of propane and/or propene from theabsorbent and the oxidation to acrolein and/or acrylic acid.

In the present invention, the gas B may also be a gas mixture.

In step (a), gas mixtures A comprising any desired amounts of propaneand/or propene can be used. Preferably, the gas mixture A containspropane and propene in a molar ratio of from 0:100 to 100:0, inparticular from 10:90 to 90:10, frequently from 80:20 to 40:60.

Preferably, the gas mixture A contains at least one further componentwhich differs from propane and/or propene and is not subject to anyparticular restrictions. As a rule, the further components depend on theorigin of the gas mixture. In particular, they comprise at least onecomponent selected from nitrogen, hydrogen, oxides of carbon, such ascarbon monoxide or carbon dioxide, further secondary componentsoriginating from a propane dehydrogenation, secondary componentsoriginating from a gas-phase oxidation of propene to acrolein and/oracrylic acid or secondary components originating from an oxidation ofpropane to acrolein and/or acrylic acid. Frequently, at least hydrogen,nitrogen, an oxide of carbon or a mixture of these is present as afurther component.

Suitable absorbents in step (a) are in principle all absorbents whichare capable of absorbing propane and/or propene. The absorbent ispreferably an organic solvent which is preferably hydrophobic and/orhigh-boiling. Advantageously, this solvent has a boiling point (at anatmospheric pressure of 1 atm) of at least 120° C., preferably of atleast 180° C., especially from 200 to 350° C., in particular from 250 to300° C., more preferably from 260 to 290° C. Expediently, the flashpoint(at an atmospheric pressure of 1 atm) is above 110° C. In general,suitable absorbents are relatively nonpolar organic solvents, forexample aliphatic hydrocarbons which preferably contain no externallyacting polar group, but also aromatic hydrocarbons. In general, it isdesirable for the absorbent to have a very high boiling point incombination with very high solubility for propane and/or propene.Examples of absorbents are aliphatic hydrocarbons, for exampleC₈-C₂₀-alkanes or C₈-C₂₀-alkenes, or aromatic hydrocarbons, for examplemiddle oil fractions from paraffin distillation or ethers having bulkygroups on the oxygen atom, or mixtures thereof, it being possible to addto them a polar solvent, for example, the 1,2-dimethyl phthalatedisclosed in DE-A 43 08 087. Esters of benzoic acid and phthalic acidwith straight-chain alkanols of 1 to 8 carbon atoms, such as n-butylbenzoate, methyl benzoate, ethyl benzoate, dimethyl phthalate anddiethyl phthalate, and thermal oils, such as biphenyl, diphenyl etherand mixtures of biphenyl and diphenyl ether or their chlorinederivatives and triarylalkenes, for example4-methyl-4′-benzyldiphenylmethane and its isomers2-methyl-2′-benzyldiphenylmethane, 2-methyl-4′-benzyldiphenylmethane and4-methyl-2′ -benzyldiphenylmethane, and mixtures of such isomers arefurthermore suitable. A suitable absorbent is a solvent mixturecomprising biphenyl and diphenyl ether, preferably in the azeotropiccomposition, in particular comprising about 25% by weight of biphenyland about 75% by weight of diphenyl ether, for example the commerciallyavailable diphyl. Frequently, this solvent mixture contains an addedsolvent, such as dimethyl phthalate, in an amount of from 0.1 to 25% byweight, based on the total solvent mixture. Other particularly suitableabsorbents are octanes, nonanes, decanes, undecanes, dodecanes,tridecanes, tetradecanes, pentadecanes, hexadecanes, heptadecanes andoctadecanes, tetradecane having proven particularly suitable. It isadvantageous if the absorbent used on the one hand has theabovementioned boiling point but on the other hand simultaneously hasnot too high a molecular weight. Advantageously, the molecular weight ofthe absorbent is ≦300 g/mol. The liquid paraffins of 8 to 10 carbonatoms which are described in DE-A 33 13 573 are also suitable. Examplesof suitable commercial products are products sold by Haltermann, such asHalpasols i, e.g. Halpasol 250/340 i and Halpasol 250/275 i, andprinting ink oils with the names PKWF and Printosol.

The absorption procedure is not subject to any particular restrictions.All processes and conditions familiar to a person skilled in the art maybe used. Preferably, the gas mixture is brought into contact with theabsorbent at from 1 to 50, preferably from 2 to 20, more preferably from5 to 10, bar and from 0 to 100° C., in particular from 30 to 50° C. Theabsorption can be carried out both in columns and in quench apparatuses.The cocurrent or the countercurrent procedure may be employed. Suitableabsorption columns are, for example, tray columns (having bubble traysand/or sieve trays), columns having stacked packings (for example sheetmetal packings having a specific surface area of from 100 to 500 m²/m³,for example Mellapak® 250 Y) and columns having dumped packings (forexample filled with Raschig packings). However, trickle and spraytowers, graphite block absorbers, surface absorbers, such as thick-filmand thin-film absorbers, and plate scrubbers, cross-spray scrubbers androtary scrubbers may also be used. It may also be advantageous to allowthe absorption to take place in a bubble column with or withoutinternals.

The separation of the propane and/or propene from the absorbent can beeffected by stripping, flashing and/or distillation.

The separation of the propane and/or propene from the absorbent in step(b) is preferably effected by stripping or desorption with a gas whichis inert with respect to the novel step (c) and/or with molecular oxygen(for example air). Here, the stripping can be carried out in the usualmanner via a pressure and/or temperature change, preferably at from 0.1to 10, in particular from 1 to 5, more preferably from 1 to 2, bar andfrom 0 to 200° C., in particular from 20 to 100° C., more preferablyfrom 30 to 50° C. Another gas suitable for the stripping is, for examplesteam, but oxygen/nitrogen mixtures are particularly preferred, forexample air. With the use of air or oxygen/nitrogen mixtures in whichthe oxygen content is more than 10% by volume, it may be expedient toadd, before or during the stripping process, a gas which reduces theexplosion range. Particularly suitable for this purpose are gases havinga specific heat capacity of >29 J/mol·K at 20° C., for example methane,ethane, propane, butane, pentane, hexane, benzene, methanol, ethanol,ammonia, carbon dioxide or water. Bubble columns with or withoutinternals are also particularly suitable for the stripping.

The separation of the propane and/or propene from the absorbent can alsobe effected by means of distillation, it being possible to use thecolumns familiar to a person skilled in the art and containing stackedpackings, dumped packings or corresponding internals.

Preferred conditions during the distillation are a pressure of from 0.01to 5, in particular from 0.1 to 3, more preferably from 1 to 2, bar anda temperature (at the bottom) of from 50 to 300° C., in particular from150 to 250° C.

Where the gas mixture A contains water, the absorption is advantageouslycombined with a condensation of the water (i.e. water quench). It isalso advantageous to follow the desorption step with a water quench inorder to minimize the losses of absorbent.

Frequently, step (c) is carried out directly after step (b), i.e.without process steps in between or intermediate stages. However, afterstep (b) and before step (c) absorbent can be separated, for example bya water quench.

The oxidation of propene and/or propane to acrolein and/or acrylic acid,which is carried out in step (c), can be effected according to allprocesses which are known to a person skilled in the art and are notsubject to any restrictions. In step (c), a one-stage or two-stageoxidation of propene to acrolein and/or acrylic acid or an oxidation ofpropane to acrolein and/or acrylic acid or both, i.e. simultaneousoxidation of propane and propene to acrolein and/or acrylic acid, can becarried out. The oxidation is expediently carried out as a selective,heterogeneously catalyzed gas-phase oxidation with molecular oxygen togive an acrolein- and/or acrylic acid-containing product gas mixture. Ifrequired, the propane and/or propene fed to the oxidation is broughtbeforehand, by indirect heat exchange, to the reaction temperaturerequired for the oxidation reaction.

In a preferred embodiment, step (c) of the novel process is carried outas an oxidation of propene to acrolein and/or acrylic acid.

In principle, the heterogeneously catalyzed gas-phase oxidation ofpropene to acrylic acid with molecular oxygen takes place in two stepsin succession along the reaction coordinate, the first of which leads toacrolein and the second from acrolein to acrylic acid. This reactionsequence in two successive steps makes it possible, in a manner knownper se, to carry out the step (c) of the novel process in two oxidationzones arranged one behind the other, it being possible for the oxidiccatalyst to be used to be adapted in an optimum manner in each of thetwo oxidation zones. Thus, as a rule a catalyst based on multimetaloxides containing the combination of elements Mo—Bi—Fe is preferred forthe first oxidation zone (propene→acrolein), while catalysts based onmultimetal oxides containing the combination of elements Mo—B areusually preferred for the second oxidation zone (acrolein→acrylic acid).Corresponding multimetal oxide catalysts for the two oxidation zoneshave been widely described and are well known to a person skilled in theart. For example, EP-A-0 253 409 refers to corresponding U.S. patents onpage 5. Advantageous catalysts for the two oxidation zones are alsodisclosed in DE-A 44 31 957 and DE-A 44 31 949. This applies inparticular to those of the formula I in the two abovementionedpublications. As a rule, the product mixture from the first oxidationzone is transferred without intermediate treatment into the secondoxidation zone.

The simplest form for realizing the two oxidation zones is therefore atube-bundle reactor, within which the catalyst load changescorrespondingly along the individual catalyst tubes with the end of thefirst reaction step. Such oxidations are described, for example, inEP-A-0 911 313, EP-A-0 979 813, EP-A-0 990 636 and DE-A 28 30 765). Ifrequired, the catalyst load in the catalyst tubes is interrupted by aninert bed.

Preferably, however, the two oxidation zones are realized in the form oftwo tube-bundle systems connected in series. These may be present in areactor, the transition from one tube bundle to another tube bundlebeing formed by a bed of inert material not housed in the catalyst tubeand expediently accessible. While the heat-transfer medium generallyflows around the catalyst tubes, said heat-transfer medium does notreach an inert bed installed as described above. The two catalyst tubebundles are therefore advantageously housed in reactors spatiallyseparated from one another. As a rule, an intermediate condenser ispresent between the two tube-bundle reactors in order to reduce anysubsequent acrolein combustion in the product gas mixture which leavesthe first oxidation zone. Instead of tube-bundle reactors, plate-typeheat exchanger reactors with salt cooling and/or evaporative cooling, asdescribed, for example, in DE-A 199 29 487 and DE-A 199 52 964, can alsobe used.

The reaction temperature in the first oxidation zone is as a rule from300 to 450° C., preferably from 320 to 390° C. The reaction temperaturein the second oxidation zone is as a rule from 200 to 300° C.,frequently from 220 to 290° C. The reaction pressure in both oxidationzones is expediently from 0.5 to 5, advantageously from 1 to 3, atm. Thegas loading (l(S.T.P.)/l·h) of the oxidation catalysts with reaction gasis frequently from 1500 to 2500 h⁻¹ or up to 4000 h⁻¹ in both oxidationzones. The propene loading (l(S.T.P.)/l·h) is frequently from 50 to 300h⁻¹, in particular from 100 to 200 h⁻¹.

In principle, the two oxidation zones can be designed as described, forexample, in DE-A 198 37 517, DE-A 199 10 506, DE-A 199 10 508 and DE-A198 37 519. Usually, the external heating in the two oxidation zones, ifdesired in multizone reactor systems, is adapted in a manner known perse to the specific composition of the reaction gas mixture and to thecatalyst load.

According to the invention, it is advantageous if, in the novel process,propene-accompanying propane acts as an advantageous inert diluent gasin a heterogeneously catalyzed propene oxidation.

The total molecular oxygen required for the oxidation can be addedbeforehand to the gas B in its total amount. However, oxygen may also beadded to said gas after the first oxidation zone.

Preferably, a molar propene:molecular oxygen ratio of from 1:1 to 1:3,frequently from 1:1.5 to 1:2, is established in the first oxidationzone. In the second oxidation zone, a molar acrolein: molecular oxygenratio of from 1:0.5 to 1:2 is preferably established.

In both oxidation zones, an excess of molecular oxygen generally has anadvantageous effect on the kinetics of the gas-phase oxidation. Sincethe heterogeneously catalyzed gas-phase oxidation of the propene toacrylic acid is subject to kinetic control, the propene can in principletherefore be initially taken in a molar excess relative to the molecularoxygen, for example also in the first oxidation zone. In this case, theexcess propene factually plays the role of a diluent gas.

In principle, the heterogeneously catalyzed gas-phase oxidation ofpropene to acrylic acid can however also be realized in a singleoxidation zone. In this case, the two reaction steps take place in anoxidation reactor which is loaded with a catalyst which is capable ofcatalyzing the reaction of both reaction steps. Here, the catalyst loadcan also change continuously or abruptly along the reaction coordinatewithin the oxidation zone. In an embodiment of step (c) in the form oftwo oxidation zones connected in series, it is also possible partly orcompletely to separate, from the product gas mixture leaving the firstoxidation zone, oxides of carbon and steam contained in said mixture andformed as byproduct in the first oxidation zone, if required beforefurther passage into the second oxidation zone. Preferably, a procedurewhich does not require such a separation is chosen.

Suitable sources for the molecular oxygen required in the oxidation step(c) are both pure molecular oxygen and molecular oxygen diluted withinert gas, such as carbon dioxide, carbon monoxide, noble gases,nitrogen and/or saturated hydrocarbons.

In an expedient manner, air is used as an oxygen source at least forcovering part of the requirement of molecular oxygen. The gas B fed tothe oxidation step (c) of the novel process advantageously comprisessubstantially only propane and propene, and exclusively air is used as asource of molecular oxygen for the oxidation. If required, cooling ofthe gas B fed to step (c) can also be effected in a direct manner bymetering in cold air.

If acrolein is the desired product, the second oxidation zone isexpediently not used in step (c).

The oxidation of propene to acrolein and/or acrylic acid in step (c) canalso be carried out as described in EP-A-0 117 146, U.S. Pat. No.5,198,578 and U.S. Pat. No. 5,183,936 or analogously to DE-A 33 13 573,CA-A-1 217 502, U.S. Pat. No. 3,161,670, U.S. Pat. No. 4,532,365 and WO97/36849. Suitable processes are also described in EP-A-0 293 224,EP-A-0 253 409, DE-A 44 31 957, DE 195 08 532 or DE-A 41 32 263,particularly preferred processes being those which operate with diluentgases in the oxidation.

The oxidation of acrolein to acrylic acid can be carried out asdescribed in WO 00/39065, by means of a fluidized-bed reactor.

The oxidation of propene to acrolein and/or acrylic acid can also becarried out using the plate-type heat exchanger reactors described inDE-A 199 52 964.

In a further preferred embodiment, step (c) of the novel process iscarried out as an oxidation of propane to acrolein and/or acrylic acid.In this oxidation, propane is converted over a suitable catalyst, in oneor more stages, to acrolein and/or acrylic acid. All processes known toa person skilled in the art are suitable for this purpose. A suitableprocess is described, for example, in JP-A-10 36 311.

Catalysts suitable for the heterogeneously catalyzed gas-phase oxidationof propane to acrolein and/or acrylic acid are multimetal oxidematerials of the formula (I)MoV_(b)M¹ _(c)M² _(d)O_(n)  (I)where

-   -   M¹ is Te and/or Sb,    -   M² is at least one of the elements from the group consisting of        Nb, Ta, W, Ti, Al, Zr, Cr, Mn, Ga, Fe, Ru, Co, Rh, Ni, Pd, Pt,        La, Bi, B, Ce, Sn, Zn, Si and In,    -   b is from 0.01 to 1,    -   c is from >0 to 1, preferably from 0.01 to 1,    -   d is from >0 to 1, preferably from 0.01 to 1, and    -   n is a number which is determined by the valency and frequency        of the elements other than oxygen in (I).

Multimetal oxide materials which have stoichiometry corresponding to theformula (I) are known, cf for example EP-A-0 608 838, EP-A-0 529 853,JP-A 7-232 071, JP-A 10-57813, JP-A 2000-37623, JP-A 10-36311, WO00/29105, Proceedings ISO 99, Rimini (Italy), Sep. 10-11, 1999, G. Centiand S. Perathoner Ed., SCI Pub. 1999, EP-A-0 767 164, Catalysis Today 49(1999), 141-153, EP-A-0 962 253, Applied Catalysis A: General 194 to 195(2000), 479 to 485, JP-A 11/169716, EP-A-0 895 809, DE-A 198 35 247,JP-A 857319, JP-A 10-28862, JP-A 11-43314, JP-A 11-57479, WO 00/29106,JP-A 10-330343, JP-A 11-285637, JP-A 10-310539, JP-A 11-42434, JP-A11-343261, JP-A 11-343262, WO 99/03825, JP-A 7-53448, JP-A 2000-51693and JP-A 11-263745.

The multimetal oxides (I), (II) and (III) described below areparticularly suitable.

In the multimetal oxide materials (I) of the formula (I), M¹ is Teand/or Sb; M² is at least one of the elements from the group consistingof Nb, Ta, W, Ti, Al, Zr, Cr, Mn, Fe, Ru, Co, Rh, Ni, Pd, Pt, Bi, B andCe; b is from 0.01 to 1; c is from 0.01 to 1; d is from 0.01 to 1; and nis a number which is determined by the valency and frequency of theelements other than oxygen in (I).

The preparation of a multimetal oxide material (I) is preferably carriedout by a procedure in which a mixture of sources of the elementalconstituents of the multimetal oxide material (I) is subjected tohydrothermal treatment and the freshly forming solid is separated offand is converted into an active oxide by thermal treatment. In themultimetal oxide material (I), M¹ is preferably Te, M² is preferably Nb,b is preferably 0.1-0.6, c is preferably 0.05-0.4 and d is preferably0.01-0.6. The thermal treatment is preferably carried out at from 350 to700° C., the thermal treatment initially being effected in particular atfrom 150 to 400° C. under an oxygen-containing atmosphere and then atfrom 350 to 700° C. under an inert gas atmosphere. Suitablestoichiometries for the multimetal oxide materials (I) are those whichare disclosed in EP-A-0 608 838, WO 00/29106, JP-A 11/169716 and EP-A-0962 253.

The hydrothermal preparation of multimetal oxide active materialprecursors is familiar to a person skilled in the art (cf. for exampleApplied Catalysis A: 194 to 195 (2000), 479-485, Kinetics and Catalysis,40, No. 3 (1999), 401 to 404, Chem. Commun. (1999), 517 to 518, JP-A6/227819 and JP-A 2000/26123).

What is understood here in particular is the thermal treatment of apreferably intimate mixture of sources of the elemental constituents ofthe desired multimetal oxide material (I) in a high-pressure vessel(autoclave) in the presence of steam at superatmospheric pressure,usually at from >100 to 600° C. The pressure range is typically up to500, preferably up to 250, atm. Temperatures above 600° C. and steampressures above 500 atm may also be used, but this is not very expedientin terms of application technology. The hydrothermal treatment isparticularly advantageously carried out under conditions under whichsteam and liquid water coexist. This is possible in a temperature rangefrom >100 to 374.15° C. (critical temperature of water) using thecorresponding pressures. The amounts of water are expediently such thatthe liquid phase is capable of holding the total amount of the startingcompounds in suspension and/or solution.

However, a procedure in which the intimate mixture of the startingcompounds completely absorbs the amount of liquid water present inequilibrium with the steam is also possible.

Advantageously, the hydrothermal treatment is carried out at from >100to 300° C., preferably from 150 to 250° C. (for example from 160 to 180°C.). Based on the sum of water and sources of the elemental constituentsof the desired multimetal oxide material (I), the amount of the latterin the autoclave is as a rule at least 1% by weight. Usually, theabovementioned amount is not above 90% by weight. Amounts of from 3 to60 or from 5 to 30, frequently from 5 to 15, % by weight are typical.

During the hydrothermal treatment, stirring may or may not be effected.Particularly suitable starting compounds (sources) for the hydrothermalpreparation variant are all those which are capable of forming oxidesand/or hydroxides on heating with water under superatmospheric pressure.Oxides and/or hydroxides of the elemental constituents may also beconcomitantly or exclusively used as starting compounds for thehydrothermal preparation. As a rule, the sources are used in finelydivided form.

Suitable sources for the elemental constituents are all those which arecapable of forming oxides and/or hydroxides on heating (if necessary inair). Oxides and/or hydroxides of the elemental constituents may beconcomitantly or exclusively used as such starting compounds.

Suitable sources of the element Mo are, for example, molybdenum oxides,such as molybdenum trioxide, molybdates, such as ammonium heptamolybdatetetrahydrate and molybdenum halides, such as molybdenum chloride.

Suitable starting compounds to be used concomitantly with the element Vare, for example, vanadyl acetylacetonate, vanadates, such as ammoniummetavanadate, vanadium oxides, such as vanadium pentoxide (V₂O₅),vanadium halides, such as vanadium tetrachloride (VCl₄), and vanadiumoxyhalides, such as VOCl₃. Expediently, vanadium starting compoundswhich are concomitantly used are those which contain the vanadium inoxidation state +4.

Suitable sources for the element tellurium are tellurium oxides, such astellurium dioxide, metallic tellurium, tellurium halides, such as TeCl₂,and telluric acids, such as orthotelluric acid H₆TeO_(6.)

Advantageous antimony starting compounds are antimony halides, such asSbCl₃, antimony oxides, such as antimony trioxide (Sb₂O₃), antimonicacids, such as HSb(OH)₆, and antimony oxide salts, such as antimonyoxide sulfate (SbO₂)SO_(4.)

Suitable niobium sources are, for example, niobium oxides, such asniobium pentoxide (Nb₂O₅), niobium oxyhalides, such as NbOCl₃, niobiumhalides, such as NbCl₅, and complex compounds of niobium and organiccarboxylic acids and/or dicarboxylic acids, for example oxalates andalcoholates. The Nb-containing solutions used in EP-A-0 895 809 are alsosuitable as a niobium source.

Regarding all other possible elements M², particularly suitable startingcompounds are their halides, nitrates, formates, oxalates, acetates,carbonates and/or hydroxides. Suitable starting compounds are often alsotheir oxo compounds, for example tungstates, or the acids derived fromthese. Frequently, ammonium salts are also used as starting compounds.

Other suitable starting compounds are polyanions of the Anderson type,as have been described, for example, in Polyhedron 6, No. 2 (1987),213-218, and have been used, for example, in Applied Catalysis A:General 194-195 (2000), 479-485, for the preparation of suitablemultimetal oxides (I) or are disclosed in the secondary literature citedtherein. A further suitable literature source for polyanions of theAnderson type is Kinetics and Catalysis, 40 (1999), 401 to 404.

Further polyanions suitable as starting compounds are, for example,those of the Dawson or Keggin type. Preferably used starting compoundsare those which are converted into their oxides at elevated temperatureseither in the presence or in the absence of oxygen, possibly withliberation of gaseous compounds.

The hydrothermal treatment itself takes, as a rule, a period of from afew hours to a few days. A period of 48 hours is typical. It isexpedient in terms of application technology if the autoclave to be usedfor the hydrothermal treatment is coated on the inside with Teflon.Before the hydrothermal treatment, the autoclave, if required includingthe aqueous mixture contained, may be evacuated. It can then be filledwith inert gas (N₂; noble gas) before the temperature is increased. Bothmeasures may also be omitted. The aqaeous mixture may additionally beflushed with inert gas for blanketing prior to the hydrothermaltreatment. It is also expedient in terms of application technology ifthe abovementioned inert gases are used for establishingsuperatmospheric pressure in the autoclave before the hydrothermaltreatment.

The required treatment of the solid freshly formed in the course of thehydrothermal treatment and separated off after the end of thehydrothermal treatment (after the end of the hydrothermal treatment, theautoclave can be either quenched to room temperature or brought to roomtemperature slowly, i.e. over a relatively long period (for example byleaving it to stand)) is expediently carried out at from 350 to 700° C.,frequently from 400 to 650° C. or from 400 to 600° C. It can be effectedunder an oxidizing, reducing or inert atmosphere. A suitable oxidizingatmosphere is, for example, air, air enriched with molecular oxygen orair depleted in oxygen. Preferably, the thermal treatment is carried outunder an inert atmosphere, i.e. for example under molecular nitrogenand/or noble gas. Of course, the thermal treatment can also be effectedunder reduced pressure.

If the thermal treatment is carried out under a gaseous atmosphere, thismay be either stationary or flowing.

In general, the thermal treatment may take up to 24 hours or more.

The thermal treatment is preferably carried out initially under anoxidizing (oxygen-containing) atmosphere (for example under air) at from150 to 400° C. or from 250 to 350° C. The thermal treatment is thenexpediently continued under inert gas at from 350 to 700° C. or from 400to 650° C. or from 400 to 600° C. The thermal treatment of thehydrothermally produced catalyst precursor can also be effected in sucha way that the catalyst precursor material is first pelleted, thenthermally treated and subsequently converted into chips.

It is expedient in terms of application technology if the solid obtainedin the hydrothermal treatment is converted into chips for the subsequentthermal treatment.

If the starting compounds used for the preparation of the multimetaloxide materials (I) are the same as those used for a conventionalpreparation of multimetal oxides (I) and the thermal treatment of theconventionally produced intimate dry blend is carried out in the sameway as the thermal treatment of the hydrothermally obtained solid, themultimetal oxide materials (I) generally have a higher selectivity ofthe acrylic acid formation and a higher activity with respect to theheterogeneously catalyzed gas-phase oxidation of propane to acrylic acidunder the same conditions.

The multimetal oxide materials (I) can be used as such (for exampleafter comminution to a powder or to chips) or can be converted intomoldings before being used. The catalyst bed may be either a fixed bed,a moving bed or a fluidized bed.

The X-ray diffraction pattern of the multimetal oxide materials (I)corresponds as a rule essentially to those in EP-A-0 529 853, DE-A 19835 247 and EP-A-0 608 838.

The multimetal oxide materials (I) can also be used in a form dilutedwith finely divided, for example colloidal, materials, such as silica,titanium dioxide, alumina, zirconium oxide or niobium oxide.

The mass dilution ratio may be up to 9 (diluent): 1 (active material).This means that possible mass dilution ratios are also 6 (diluent): 1(active material) and 3 (diluent): 1 (active material). The diluent canbe incorporated before and/or after the calcination. As a rule, thediluent is incorporated before the hydrothermal treatment. If theincorporation is effected before the calcination, the diluent must bechosen so that it is substantially retained as such during thecalcination. The same applies to the hydrothermal treatment in the caseof incorporation before said treatment is carried out. As a rule, thisis true, for example, in the case of oxides calcined at correspondinglyhigh temperatures.

Other catalysts suitable for the propane oxidation are multimetal oxidematerials (II) which have the abovementioned formula (I) and whose X-raydiffraction pattern has reflections h, i and k whose peaks are at thediffraction angles (2θ) 22.2±0.4° (h), 27.3±0.4° (i) and 28.2±0.4° (k),where

-   -   the reflection h is the one with the highest intensity within        the X-ray diffraction pattern and has a half-width of not more        than 0.5°,    -   the intensity Pi of the reflection i and the intensity P_(k) of        the reflection k fulfill the relationship 0.65≦R≦0.85, in which        R is the intensity ratio defined by the formula        R=P_(i)/(P_(i)+P_(k)),        and    -   the half-width of the reflection i and of the reflection k is        ≦1° in each case.

Preferably, 0.67 <R<0.75 and very particularly preferably R=0.70 to 0.75or R 0.72.

The use of multimetal oxide materials (II) where M¹ is Te is preferred.Furthermore, those multimetal oxide materials (II) in which M² is Nb,Ta, W and/or Ti are advantageous. Preferably, M² is Nb.

The stoichiometric coefficient b of the multimetal oxide materials (II)is advantageously from 0.1 to 0.6. In a corresponding manner, thepreferred range for the stoichiometric coefficient c is from 0.01 to 1or from 0.05 to 0.4 and advantageous values for d are from 0.01 to 1 orfrom 0.1 to 0.6. Particularly advantageous multimetal oxide materials(II) are those in which the stoichiometric coefficients b, c and d aresimultaneously in the abovementioned preferred ranges. Further suitablestoichiometries for the multimetal oxide materials (II) are those whichare disclosed in the publications of the prior art cited above, inparticular those disclosed in JP-A 7-53448.

A specific process for the preparation of multimetal oxide materials(II) is disclosed, for example, in JP-A 11-43314, in which the relevantmultimetal oxide materials (II) are recommended as catalysts for theheterogeneously catalyzed oxydehydrogenation of ethane to ethene.

Thereafter, a multimetal oxide material of the formula (I) which is amixture of the i-phase and other phases (for example k-phase) is firstproduced in a manner known per se and disclosed in the cited prior artpublications. In this mixture, for example, the proportion of i-phasecan be increased by selectively removing the other phases, for examplethe k-phase, under the microscope or washing the multimetal oxidematerial with suitable liquids. Suitable such liquids are, for example,aqueous solutions of organic acids (for example oxalic acid, formicacid, acetic acid, citric acid and tartaric acid), inorganic acids (forexample nitric acid), alcohols and aqueous hydrogen peroxide solutions.Furthermore, JP-A 7-232071 discloses a process for the preparation ofmultimetal oxide materials (II).

Multimetal oxide materials (II) are obtainable by the preparation methodaccording to DE-A 198 35 247. According to this, a very intimate,preferably finely divided, dry blend is produced from suitable sourcesof their elemental constituents and said dry blend is subjected to athermal treatment at from 350 to 700° C. or from 400 to 650° C. or from400 to 600° C. The thermal treatment can be carried out under either anoxidizing, reducing or inert atmosphere. A suitable oxidizing atmosphereis, for example, air, air enriched with molecular oxygen or air depletedin oxygen. Preferably, the thermal treatment is carried out under aninert atmosphere, i.e. for example under molecular nitrogen and/or noblegas.

Usually, the thermal treatment is effected at atmospheric pressure (1atm). Of course, the thermal treatment can also be effected underreduced or superatmospheric pressure.

If the thermal treatment is carried out under a gaseous atmosphere, thismay be either stationary or flowing. In general, the thermal treatmentmay take up to 24 hours or more.

The thermal treatment is initially preferably carried out under anoxidizing (oxygen-containing) atmosphere (for example under air) at from150 to 400° C. or from 250 to 350° C. Thereafter, the thermal treatmentis expediently continued under an inert gas at from 350 to 700° C. orfrom 400 to 650° C. or from 400 to 600° C. The thermal treatment canalso be effected in such a way that the catalyst precursor material isfirst pelleted (if required after pulverization and, if required, withthe addition of from 0.5 to 2% by weight of finely divided graphite)before its thermal treatment, then subjected to the thermal treatmentand subsequently converted into chips.

The thorough mixing of the starting compounds in the preparation of themultimetal oxide materials (II) can be effected in dry or in wet form.If it is effected in dry form, the starting compounds are expedientlyused in the form of finely divided powders and, after mixing and anycompaction, are subjected to the calcination (thermal treatment).However, the thorough mixing is preferably effected in wet form.Usually, the starting compounds are mixed with one another in the formof an aqueous solution and/or suspension. Thereafter, the aqueousmaterial is dried and is calcined after the drying. Expediently, theaqueous material is an aqueous solution or an aqueous suspension.Preferably, the drying process is carried out immediately after thepreparation of the aqueous mixture and by spray-drying (the outlettemperatures are as a rule from 100 to 150° C.; the spray-drying can becarried out by the cocurrent or countercurrent method), which requires aparticularly intimate dry blend, especially when the aqueous material tobe spray-dried is an aqueous solution or suspension.

Suitable sources or starting compounds for the multimetal oxide material(II) are the compounds described above in the case of the multimetaloxide material (I).

The multimetal oxide materials (II) can be converted into moldings as inthe case of the multimetal oxide materials (I) and can be used in thesame way as these. The shaping of the multimetal oxide materials (II)can be effected, for example, by application to a support, as describedbelow under catalyst (III), or by extrusion and/or pelleting, both offinely divided multimetal oxide material (II) and of finely dividedprecursor material of a multimetal oxide material (II).

In the same way as the multimetal oxide materials (I), the multimetaloxide materials (II) can also be used in a form diluted with finelydivided materials.

Suitable geometries are spheres, solid cylinders and hollow cylinders(rings). The longest dimension of the abovementioned geometries is as arule from 1 to 10 mm. In the case of cylinders, their length ispreferably from 2 to 10 mm and their external diameter preferably from 4to 10 mm. In the case of rings, the wall thickness is moreover usuallyfrom 1 to 4 mm. Suitable annular unsupported catalysts may also have alength from 3 to 6 mm, an external diameter of from 4 to 8 mm and a wallthickness of from 1 to 2 mm. However, an annular unsupported catalysthaving the dimensions of 7 mm×3 mm×4 mm or of 5 mm×3 mm×2 mm (externaldiameter x length x internal diameter) is also possible.

The definition of the intensity of a reflection in the X-ray diffractionpattern is based here on the definition stated in DE-A 198 35 247 andthat in DE-A 100 51 419 and DE-A 100 46 672.

This means that if A¹ is the peak of a reflection 1 and B¹ is the nextpronounced minimum (minima having the reflection shoulders are not takeninto account) to the left of peak A¹ in the line of the X-raydiffraction pattern when viewed along the intensity axis perpendicularto the 2θ axis and B² is correspondingly the next pronounced minimum tothe right of the peak A¹ and C¹ is a point at which a straight linedrawn from the peak A¹ perpendicular to the 2θ axis intersects astraight line connecting the points B¹ and B², then the intensity ofreflection 1 is the length of the straight line section A¹C¹ whichextends from the peak A¹ to the point C¹. The expression minimum means apoint at which the slope of a tangent to the curve in a base region ofreflection 1 changes from a negative value to a positive value, or apoint at which the slope tends to zero, the coordinates of the 2θ axisand of the intensity axis being used for specifying the slope.

Here, the half-width is correspondingly the length of the straight linesection which is present between the two points of intersection H¹ andH² if a parallel to the 2θ axis is drawn in the center of the straightline section A¹C¹, H¹ and H² being in each case the first point ofintersection of this parallel, to the left and right of A¹, with thatline of the X-ray diffraction pattern which is defined as above.

An exemplary procedure for the determination of half-width and intensityis also shown in FIG. 6 in DE-A 100 46 672.

In addition to the reflections h, i and k, the X-ray diffraction patternof advantageous catalytically active multimetal oxide materials (II)contains, as a rule, further reflections whose peaks are at thefollowing diffraction angles (2θ):

-   -   9.0±0.4° (l),    -   6.7±0.4° (o) and    -   7.9±0.4° (p).

It is advantageous if the X-ray diffraction pattern of the catalyticallyactive oxide materials of the formula (I) additionally contains areflection whose peak is at the following diffraction angle (2θ):

-   -   45.2±0.4° (q).

Frequently, the X-ray diffraction pattern of the multimetal oxidematerials (II) also contains the reflections 29.2±0.4° (m) and 35.4±0.4°(n).

The multimetal oxide material (II) may be one whose X-ray diffractionpattern has no reflection with a peak position of 2θ=50.0±0.3°, i.e. onewhich contains no k-phase.

However, the multimetal oxide material (II) can also contain a k-phase,its X-ray diffraction pattern generally also containing furtherreflections whose peaks are at the following diffraction angles (2θ):

-   -   36.2±0.40 and    -   50.0±10.4°.

If the reflection h is assigned the intensity 100, it is advantageous ifthe reflections i, l, m, n, o, p and q have the following intensities onthe same intensity scale:

-   -   i: from 5 to 95, frequently from 5 to 80, in some cases from 10        to 60;    -   l: from 1 to 30;    -   m: from 1 to 40;    -   n: from 1 to 40;    -   o: from 1 to 30;    -   p: from 1 to 30 and    -   q: from 5 to 60.

If the X-ray diffraction pattern contains additional reflections fromthose stated above, the half-width thereof is as a rule <1°.

All data based here on an X-ray diffraction pattern relate to an X-raydiffraction pattern produced using Cu—Kα radiation as the X-radiation(Siemens diffractometer Theta-Theta D-5000, tube voltage: 40 kV, tubecurrent: 40 mA, aperture V20 (variable), collimator V20 (variable),secondary monochromator aperture (0.1 mm), detector aperture (0.6 mm),measuring interval (2θ): 0.02°, measuring time per step: 2.4 s,detector: scintillation counter).

The specific surface area of multimetal oxide materials (II) is oftenfrom 1 to 30 m²/g (BET surface area, nitrogen).

Another suitable catalyst for the propane oxidation is a catalyst (III)which consists of a support and a catalytically active oxide material ofthe abovementioned formula (I) which is applied to the surface of thesupport.

The use of oxide materials of the formula (I) where M¹ is Te ispreferred. It is furthermore advantageous if M² is Nb, Ta, W and/or Ti.Preferably, M² is Nb.

The stoichiometric coefficient b of the oxide materials of the formula(I) in catalyst (III) is advantageously from 0.1 to 0.6. In acorresponding manner, the preferred range for the stoichiometriccoefficient c is from 0.01 to 1 or from 0.05 to 0.4, and advantageousvalues for d are from 0.01 to 1 or from 0.1 to 0.6. Particularlyadvantageous oxide materials are those in which the stoichiometriccoefficients b, c and d are simultaneously in the abovementionedpreferred ranges.

Further suitable stoichiometries for the oxide materials of the formula(I) are those which are disclosed in the abovementioned publications, inparticular those which are disclosed in EP-A-0 608 838, WO 00/29106,JP-A 11/169716 and EP-A-0 962 253.

The application of the above-described multimetal oxide material (II) asthe oxide material of the formula (I) to a support is also particularlypreferred for the preparation of the catalyst (III).

The supports are preferably chemically inert, i.e. they playsubstantially no part in the course of the catalytic gas-phase oxidationof propane to acrylic acid. Particularly suitable materials for thesupports are alumina, silica, silicates, such as clay, kaolin, steatite,pumice, aluminum silicate and magnesium silicate, silicon carbide,zirconium dioxide and thorium dioxide.

The surface of the support may be either smooth or rough.Advantageously, the surface of the support is rough since increasedsurface roughness generally results in high adhesive strength of theapplied active material coat.

Frequently, the surface roughness Rz of the support is from 5 to 200 μm,often from 20 to 100 μm (determined according to DIN 4768, Sheet 1,using a Hommel tester for DIN-ISO measured surface variables fromHommelwerke, Germany).

Furthermore, the support material may be porous or nonporous. Thesupport material is expediently nonporous (total volume of the pores ≦1%by volume, based on the volume of the support).

The thickness of the active oxide material coat present on the coatedcatalyst is usually from 10 to 1000 mm. However, it may also be from 50to 700 μm, from 100 to 600 μm or from 300 to 500 μm or from 150 to 400μm. Possible coated thicknesses are also from 10 to 500 μm, from 100 to500 μm or from 150 to 300 μm.

In principle, any desired geometries of the supports are suitable. Theirlongest dimension is as a rule from 1 to 10 mm. However, spheres orcylinders, in particular hollow cylinders (rings), are preferably usedas supports. Advantageous diameters of the support spheres are from 1.5to 4 mm. If cylinders are used as supports, their length is preferablyfrom 2 to 10 mm and their external diameter preferably from 4 to 10 mm.In the case of rings, the wall thickness is moreover usually from 1 to 4mm. Suitable annular supports can also have a length of from 3 to 6 mm,an external diameter of from 4 to 8 mm and a wall thickness of from 1 to2 mm. However, an annular support having measurements of 7 mm×3 mm×4 mmor of 5 mm×3 mm×2 mm (external diameter x length x internal diameter) isalso possible.

The preparation of the catalysts (III) can be carried out in a verysimple manner by preforming active oxide materials of the formula (I),converting them into a finely divided form and finally applying them tothe surface of the support with the aid of a liquid binder. For thispurpose, the surface of the support is moistened in a very simple mannerwith the liquid binder and a coat of the active material is caused toadhere to the moistened surface by bringing said surface into contactwith finely divided active oxide materials of the formula (I). Finally,the coated support is dried. Of course, the process can be repeatedperiodically to achieve a thicker coat. In this case, the coated parentstructure becomes the new support, etc.

The fineness of the catalytically active oxide material of the formula(I) which is to be applied to the surface of the support is adapted tothe desired coat thickness. For example, those active material powdersof which at least 50% of the total number of powder particles passthrough a sieve of mesh size of from 1 to 20 μm and whose numericalproportion of particles having a longest dimension above 50 μm is lessthan 10% are suitable for the coat thickness range of from 100 to 500μm. As a rule, the distribution of the longest dimensions of the powderparticles corresponds to a Gaussian distribution as a result of thepreparation.

For carrying out the described coating process on an industrial scale,it is advisable, for example, to use the process principle disclosed inDE-A 29 096 71. There, the supports to be coated are initially taken ina preferably inclined (the angle of inclination is as a rule ≧0° and≦90°, in general ≧30° and ≦90°; the angle of inclination is the angle ofthe central axis of the rotating container relative to the horizontal)rotating container (for example rotating pan or coating drum). Therotating container transports the supports, which for example arespherical or cylindrical, under two metering apparatuses arranged aspecific distance apart. The first of the two metering apparatusesexpediently corresponds to a nozzle (for example an atomizer nozzleoperated with compressed air) through which the supports rolling in therotating pan are sprayed and moistened in a controlled manner with theliquid binder. The second metering apparatus is located outside theatomization cone of the liquid binder sprayed in and serves for feedingin the finely divided oxidic active material (for example via a shakingconveyor or a powder screw). The support spheres moistened in acontrolled manner take up the supplied active material powder which, asa result of the rolling movement, becomes compacted to a cohesive coaton the outer surface of the support, which for example is cylindrical orspherical.

If required, the support provided with the basecoat in this mannerpasses through the spray nozzles again in the course of the subsequentrevolution, is moistened in a controlled manner in order to be able totake up a further coat of finely divided oxidic active material in thecourse of the further movement, etc. (intermediate drying is as a rulenot necessary). Finely divided oxidic acid material and liquid binderare as a rule fed in continuously and simultaneously.

The liquid binder can be removed after the end of the coating, forexample by the action of hot gases, such as N₂ or air. The coatingprocess described is known to provide satisfactory adhesion both of thesuccessive coats to one another and of the basecoat to the surface ofthe support.

What is important for the coating procedure described above is that themoistening of the support surface to be coated is carried out in acontrolled manner. Briefly, this means that the support surface isexpediently moistened so that it has adsorbed liquid binder but noliquid phase as such is visible on the support surface. If the supportsurface is too moist, the finely divided catalytically active oxidematerial forms separate agglomerates instead of being deposited on thesurface. Detailed information in this context is to be found in DE-A 2909 671.

The abovementioned final removal of the liquid binder used can beeffected in a controlled manner, for example by evaporation and/orsublimation. In the simplest case, this can be effected by the action ofhot gases of corresponding temperature (frequently from 50 to 300° C.,often 150° C.). However, only preliminary drying can be effected by theaction of hot gases. The final drying can then be carried out, forexample, in a drying oven of any desired type (for example a belt dryer)or in the reactor. The temperature employed should not be above thecalcination temperature used for the preparation of the oxidic activematerial. Of course, the drying can also be carried out exclusively in adrying oven.

Regardless of type and geometry of the support, the following can beused as binders for the coating process: water, monohydric alcohols,such as ethanol, methanol, propanol and butanol, polyhydric alcohols,such as ethylene glycol, 1,4-butanediol, 1,6-hexanediol or glycerol,monobasic or polybasic organic carboxylic acids, such as propionic acid,oxalic acid, malonic acid, glutaric acid or maleic acid, aminoalcohols,such as ethanolamine or diethanolamine, and monofunctional orpolyfunctional organic amides, such as formamide. Advantageous bindersare also solutions consisting of from 20 to 90% by weight of water andfrom 10 to 80% by weight of an organic compound which is dissolved inwater and whose boiling point or sublimation temperature at atmosphericpressure (1 atm) is >100° C., preferably >150° C. The organic compoundis advantageously selected from the above list of possible organicbinders. Preferably, the organic fraction of the abovementioned aqueousbinder solutions is from 10 to 50, particularly preferably from 20 to30, % by weight. Other suitable organic components are monosaccharidesand oligosaccharides, such as glucose, fructose, sucrose or lactose, andpolyethylene oxides and polyacrylates.

The preparation of the catalytically active oxide materials of theformula (I) can be carried out in a manner known per se, as in the priorart publications cited above, i.e. the preparation can be carried out,for example, both hydrothermally and in a conventional manner.

In the latter case, the catalytically active oxide materials of theformula (I) are obtainable by producing from suitable sources of theirelemental constituents a very intimate, preferably finely divided dryblend and subjecting the latter to a thermal treatment at from 350 to700° C. or from 400 to 650° C. or from 400 to 600° C. The thermaltreatment can be effected as described above in the case of themultimetal oxide material (II). The thorough mixing of the startingcompounds can also be carried out as described above in the case ofmultimetal oxide material (II).

Suitable sources of the elemental constituents when carrying out theabove-described preparation procedure for the catalytically active oxidematerials of the formula (I) are the starting compounds or sourcesdescribed above in the case of the multimetal oxide material (I).

Coated catalysts which have the multimetal oxide material (II) ascatalytically active oxide material of the formula (I) are particularlypreferred.

However, the active oxide materials of the formula (I) from WO 00/29106,which substantially have an amorphous structure which appears in theX-ray diffraction pattern in the form of very broad reflections havingpeaks at the 2θ angles of about 22° and about 27°, are also suitable forproducing the coated catalysts.

However, the active oxide materials of the formula (I) from EP-A-0 529853 and from EPA-0 608 838, which have very narrow reflections at 2θpeak positions of 22.1±0.3°, 28.2±0.3°, 36.2±0.3°, 45.2±0.3° and50.0±0.3° in the X-ray diffraction pattern, are also suitable.

The coated catalysts can be prepared not only by applying the finished,finely milled active oxide materials of the formula (I) to the moistenedsupport surface; instead of the active oxide material, a finely dividedprecursor material thereof can also be applied to the moistened supportsurface (using the same coating process and binder) and the calcinationcan be carried out after drying of the coated support. A suitable finelydivided precursor material of this type is, for example, the materialwhich is obtainable by first producing, from the sources of theelemental constituents of the desired active oxide material of theformula (I), a very intimate, preferably finely divided, dry blend (forexample by spray-drying an aqueous suspension or solution of thesources) and subjecting this finely divided dry blend (if necessaryafter pelleting with addition of from 0.5 to 2% by weight of finelydivided graphite) to a thermal treatment for a few hours at from 150 to350° C., preferably from 250 to 350° C. under an oxidizing(oxygen-containing) atmosphere (for example under air) and, if required,finally subjecting it to milling. After the coating of the supports withthe precursor material, calcination is then effected, preferably underan inert gas atmosphere (all other atmospheres are also suitable) atfrom 360 to 700° C. or from 400 to 650° C. or from 400 to 600° C.

The multimetal oxide materials (II) described above or the catalysts(III) comprising the multimetal oxide material (II) as the catalyticallyactive oxide material can also be used for the oxidation of propene.Here, the propene can be oxidized in the presence of propane. If propaneis used as a diluent gas, some of it can also be oxidized to acrylicacid.

The procedure for the propane oxidation is not subject to anyrestrictions. It can be carried out according to all processes known toa person skilled in the art. For example, the procedure described inEP-A-0 608 838 or WO 00/29106 can be employed, i.e. a gas B with whichthe catalyst is to be loaded at reaction temperatures of, for example,from 200 to 550° C. or from 230 to 480° C. or from 300 to 440° C. instep (c) may have, for example, the following composition:

-   -   from 1 to 15, preferably from 1 to 7, % by volume of propane,    -   from 44 to 99% by volume of air and    -   from 0 to 55% by volume of steam.

Other possible compositions of the gas mixture fed to step (c) forproducing the gas B are:

-   -   from 70 to 95% by volume of propane,    -   from 5 to 30% by volume of molecular oxygen and    -   from 0 to 25% by volume of steam.

The plate-type heat exchanger reactors described in DE-A 199 52 964 arealso suitable for carrying out the propane oxidation. In anotherembodiment of the present invention, the propane oxidation is carriedout according to the processes described in DE-A 198 37 517, DE-A 198 37518, DE-A 198 37 519 and DE-A 198 37 520.

The product gas mixture leaving the propene oxidation and/or propaneoxidation as step (c) of the novel process does not exclusively consistof the desired product acrolein and/or acrylic acid but is as a rulecomposed substantially of acrolein and/or acrylic acid, unconvertedmolecular oxygen, propane, propene, molecular nitrogen, steam formed asa byproduct and/or concomitantly used as diluent gas, oxides of carbonwhich are formed as a byproduct and/or concomitantly used as diluentgas, and small amounts of other lower aldehydes, hydrocarbons and otherinert diluent gases.

The desired product acrolein and/or acrylic acid can be separated fromthe product gas mixture in a manner known per se, for example byazeotropic separation, fractional distillation (with or without asolvent) or crystallization. For example, partial condensation of theacrylic acid, absorption of acrylic acid in water or in a high-boilinghydrophobic organic solvent or absorption of acrolein in water or inaqueous solutions of lower carboxylic acids with subsequent working-upof the absorbates is suitable; alternatively, the product gas mixturecan also be subjected to fractional condensation, cf. for example EP-A-0117 146, DE-A 43 08 087, DE-A 43 35 172, DE-A 44 36 243, DE-A 199 24 532and DE-A 199 24 533.

In a particularly preferred embodiment of the novel process, after step(c) has been carried out and the desired product isolated, unreactedpropane and/or propene are then separated from the remaining gas mixtureaccording to the invention in steps (a) and (b) and are recycled to step(c).

The gas mixture A used in step (a) of the novel process may also be agas mixture which has the composition of a gas mixture which isobtainable by catalytic dehydrogenation of propane to propene. Here, thedehydrogenation can be effected by oxidation, i.e. by supplying oxygen,or without a supply of oxygen, in particular substantially without asupply of oxygen. In the dehydrogenation with a supply of oxygen, adistinction may be made between two cases. In one case, all hydrogenformed is oxidized by an excess of oxygen so that the product gas nolonger contains any hydrogen but excess oxygen (oxidativedehydrogenation). In the second case, only sufficient oxygen is added tocover the enthalpy of reaction, so that no oxygen is contained in theproduct gas but hydrogen may well be (autothermal procedure). Thepropane dehydrogenation can be carried out catalytically orhomogeneously (noncatalytically).

Dehydrogenation of propane can be carried out, for example, as describedin DE-A 33 13 573 and EP-A-0 117 146.

In principle, the oxidative propane dehydrogenation can be carried outas a homogeneous and/or heterogeneously catalyzed oxydehydrogenation ofpropane to propene with molecular oxygen. If this first reaction stageis designed as a homogeneous oxydehydrogenation, it can be carried outin principle as described, for example, in U.S. Pat. No. 3,798,283,CN-A-1 105 352, Applied Catalysis 70(2) (1991), 175-187, Catalysis Today13 (1992), 673-678, and DE-A-196 22 331, it also being possible to useair as the oxygen source.

The temperature of the homogeneous oxydehydrogenation is expedientlychosen to be in the range from 300 to 700° C., preferably from 400 to600° C., particularly preferably from 400 to 500° C. The operatingpressure may be from 0.5 to 100, in particular from 1 to 10, bar. Theresidence time is usually from 0.1 or 0.5 to 20, preferably from 0.1 or0.5 to 5, seconds. The reactor used may be, for example, a tubularfurnace or a tube-bundle reactor, for example a countercurrent tubularfurnace with stack gas as a heat-transfer medium or tube-bundle reactorwith salt melt as a heat-transfer medium. The propane-to-oxygen ratio inthe starting mixture is preferably from 0.5:1 to 40:1, in particularfrom 1:1 to 6:1, more preferably from 2:1 to 5:1. The starting mixturemay also comprise further, substantially inert, components, such aswater, carbon dioxide, carbon monoxide, nitrogen, noble gases and/orpropene, it also being possible for these to be recycled components.Here, components recycled to stage (a) are generally referred to asrecycle gas.

If the propane dehydrogenation is designed as a heterogeneouslycatalyzed oxydehydrogenation, it can be carried out in principle asdescribed, for example, in U.S. Pat. No. 4,788,371, CN-A 1073893,Catalysis Letters 23 (1994), 103-106, W. Zhang, Gaodeng Xuexiao HuaxueXuebao 14 (1993), 566, Z. Huang, Shiyou Huagong 21 (1992), 592, WO97/36849, DE-A 197 53 817, U.S. Pat. No. 3,862,256, U.S. Pat. No.3,887,631, DE-A 195 30 454, U.S. Pat. No. 4,341,664, J. of Catalysis 167(1997), 560-569, J. of Catalysis 167 (1997), 550-559, Topics inCatalysis 3 (1996), 265-275, U.S. Pat. No. 5,086,032, Catalysis Letters10 (1991), 181-192, Ind. Eng. Chem. Res. 35 (1996), 14-18, U.S. Pat. No.4,255,284, Applied Catalysis A: General 100 (1993), 111-130, J. ofCatalysis 148 (1994), 56-67, V. Cortés Corberán and S. Vic Belló(Ed.),New Developments in Selective Oxidation II, 1994, Elsevier Science B.V.,pages 305-313, 3^(rd) World Congress on Oxidation Catalysis, R. K.Grasselli, S. T. Oyama, A. M. Gaffney and J. E. Lyons (Ed.), 1997,Elsevier Science B.V., page 375 et seq. Air may also be used as theoxygen source. Preferably, however, the oxygen source comprises at least90, more preferably 95, mol %, based on 100 mol % of the oxygen source,of oxygen.

The catalysts suitable for the heterogeneous oxydehydrogenation are notsubject to any particular restrictions. All oxydehydrogenation catalystswhich are known to a person skilled in the art in this area and whichare capable of oxidizing propane to propene are suitable. In particular,all oxydehydrogenation catalysts stated in the abovementionedpublications may be used. Preferred catalysts include, for example,oxydehydrogenation catalysts which comprise MoVNb oxides or vanadylpyrophosphate, each with a promoter. An example of a particularlysuitable catalyst is a catalyst which contains a mixed metal oxidecomprising Mo, V, Te, O and X as substantial components, where X is atleast one element selected from niobium, tantalum, tungsten, titanium,aluminum, zirconium, chromium, manganese, iron, ruthenium, cobalt,rhodium, nickel, palladium, platinum, antimony, bismuth, boron, indiumand cerium. Other particularly suitable oxydehydrogenation catalysts arethe multimetal oxide materials or multimetal oxide catalysts A ofDE-A-197 53 817, the multimetal oxide materials or multimetal oxidecatalysts A stated in the abovementioned publication as being preferredbeing very particularly advantageous. This means that particularlysuitable active materials are multimetal oxide materials (IV) of theformula IVM¹ _(a)MO_(1-b)M² _(b)O_(x)  (IV),where

-   -   M¹ is Co, Ni, Mg, Zn, Mn and/or Cu,    -   M² is W, V, Te, Nb, P, Cr, Fe, Sb, Ce, Sn and/or La,    -   a is 0.5-1.5    -   b is 0-0.5        and    -   x is a number which is determined by the valency and frequency        of the elements other than oxygen in (IV).

In principle, suitable active materials (IV) can be prepared in a simplemanner by producing, from suitable sources of their elementalconstituents, a very intimate, preferably finely divided, dry blendhaving a composition corresponding to their stoichiometry and calciningthis dry blend at from 450 to 1000° C. Suitable sources of the elementalconstituents of the multimetal oxide active materials (IV) are thosecompounds which are oxides and/or those compounds which can be convertedinto oxides by heating, at least in the presence of oxygen. These are inparticular halides, nitrates, formates, oxalates, citrates, acetates,carbonates, ammine complex salts, ammonium salts and/or hydroxides. Thethorough mixing of the starting compounds for the preparation of themultimetal oxide materials (IV) can be effected in dry form, for exampleas finely divided powder, or in wet form, for example using water as asolvent. The multimetal oxide materials (IV) can be used both in powderform and after shaping to give specific catalyst geometries, it beingpossible to carry out the shaping before or after the final calcination.Unsupported catalysts may be used, or the shaping of a pulverulentactive material or precursor material can also be effected byapplication to preshaped inert catalyst supports. Conventional, porousor nonporous aluminas, silica, thorium dioxide, zirconium dioxide,silicon carbide or silicate can be used as support materials, it beingpossible for the supports to have a regular or irregular shape.

For the heterogeneously catalyzed oxydehydrogenation of propane, thereaction temperature is preferably from 200 to 600° C., in particularfrom 250 to 500° C., more preferably from 350 to 440° C. The operatingpressure is preferably from 0.5 to 10, in particular from 1 to 10, morepreferably from 1 to 5, bar. Operating pressures above 1 bar, forexample from 1.5 to 10 bar, have proven particularly advantageous. As arule, the heterogeneously catalyzed oxydehydrogenation of propane iscarried out over a fixed catalyst bed. The latter is expediently loadedinto the tubes of a tube-bundle reactor, as described, for example, inEP-A-0 700 893 and in EP-A-0 700 714 and the literature cited in thesepublications. The average residence time of the reaction gas mixture inthe catalyst bed is expediently from 0.5 to 20 seconds. The ratio ofpropane to oxygen varies with the desired conversion and the selectivityof the catalyst and is expediently from 0.5:1 to 40:1, in particularfrom 1:1 to 6:1, more preferably from 2:1 to 5:1. As a rule, the propeneselectivity decreases with increasing propane conversion. Thepropane-to-propene reaction is therefore preferably carried out in sucha way that relatively low propane conversions are achieved incombination with high propene selectivity. The propane conversion isparticularly preferably from 5 to 40%, more preferably from 10 to 30%.Here, the term propane conversion means the proportion of suppliedpropane which is converted. The selectivity is particularly preferablyfrom 50 to 98%, more preferably from 80 to 98%, the term selectivityreferring to the number of moles of propene which are produced per moleof converted propane, expressed as a percentage.

Preferably, the starting mixture used in the oxidative propanedehydrogenation contains from 5 to 95% by weight, based on 100% byweight of starting mixture, of propane. In addition to propane andoxygen, the starting mixture for the heterogeneously catalyzedoxydehydrogenation may also comprise further, in particular inert,components, such as water, carbon dioxide, carbon monoxide, nitrogen,noble gases and/or propene. The heterogeneous oxydehydrogenation canalso be carried out in the presence of diluents, for example steam.

Any desired reactor sequence which is known to a person skilled in theart may be used for carrying out the homogeneous oxydehydrogenation orthe heterogeneously catalyzed oxydehydrogenation. For example, thereaction can be carried out in a single stage or in two or more stagesbetween which oxygen is introduced. It is also possible to usehomogeneous and heterogeneously catalyzed oxydehydrogenations incombination with one another.

As possible constituents, the product mixture of the propaneoxydehydrogenation may contain, for example, the following components:propene, propane, carbon dioxide, carbon monoxide, water, nitrogen,oxygen, ethane, ethene, methane, acrolein, acrylic acid, ethylene oxide,butane, acetic acid, formaldehyde, formic acid, propylene oxide andbutene. A preferred product mixture obtained in the propaneoxydehydrogenation contains: from 5 to 10% by weight of propene, from 1to 2% by weight of carbon monoxide, from 1 to 3% by weight of carbondioxide, from 4 to 10% by weight of water, from 0 to 1% by weight ofnitrogen, from 0.1 to 0.5% by weight of acrolein, from 0 to 1% by weightof acrylic acid, from 0.05 to 0.2% by weight of acetic acid, from 0.01to 0.05% by weight of formaldehyde, from 1 to 5% by weight of oxygen,from 0.1 to 1.0% by weight of further abovementioned components andpropane as the remainder, based in each case on 100% by weight ofproduct mixture.

In general, the propane dehydrogenation for the preparation of gasmixture A can also be carried out as a heterogeneously catalyzed propanedehydrogenation substantially in the absence of oxygen, as described inDE-A 33 13 573, or as follows.

Since the dehydrogenation reaction takes place with an increasingvolume, the conversion can be increased by reducing the partial pressureof the products. This can be achieved in a simple manner, for example bydehydrogenation at reduced pressure and/or by admixing of substantiallyinert diluent gases, for example steam, which is usually an inert gasfor the dehydrogenation reaction. Another advantage of dilution withsteam is that it generally results in reduced coking of the catalystused since the steam reacts with resulting coke according to theprinciple of gasification of coal. Moreover, steam may be present asdiluent gas in the downstream oxidation step (c). However, steam canalso easily be separated off partly or completely before step (a) (forexample by condensation), which makes it possible to increase theproportion of diluent gas N₂ when the gas mixture obtainable thereby isfurther used in oxidation step (c). Further diluents suitable for thepropane dehydrogenation are, for example, CO, CO₂, nitrogen and noblegases, such as He, Ne and Ar. All diluents stated may be present eitherby themselves or in the form of a very wide range of mixtures. It isadvantageous that said diluents are as a rule also diluents suitable forthe oxidation step (c). In general, diluents which are inert in therespective stage (i.e. which undergo chemical change to an extent ofless than 5, preferably less than 3 and more preferably less than 1, mol%) are preferred. In principle, all dehydrogenation catalysts known inthe prior art are suitable for the propane dehydrogenation. They can bedivided roughly into two groups, i.e. into those which are oxidic innature (for example chromium oxide and/or alumina) and those whichconsist of at least one, as a rule comparatively noble, metal (forexample platinum) deposited on a generally oxidic support.

Inter alia, all dehydrogenation catalysts which are recommended in WO99/46039, U.S. Pat. No. 4, 788,371, EP-A-0 705 136, WO 99/29420, U.S.Pat. No. 4,220,091, U.S. Pat. No. 5,430,220, U.S. Pat. No. 5,877,369,EP-A-0 117 146, DE-A 199 37 196, DE-A 199 37 105 and DE-A 199 37 107 canthus be used. In particular, the catalyst according to example 1,example 2, example 3 and example 4 of DE-A 199 37 107 can be used.

These are dehydrogenation catalysts which contain from 10 to 99.9% byweight of zirconium dioxide, from 0 to 60% by weight of alumina, silicaand/or titanium dioxide and from 0.1 to 10% by weight of at least oneelement of the first or second main group, one element of the thirdsubgroup, one element of the eighth subgroup of the Periodic Table ofthe Elements, lanthanum and/or tin, with the proviso that the sum of thepercentages by weight is 100% by weight.

In principle, all reactor types and process variants known in the priorart are suitable for carrying out the propane dehydrogenation.Descriptions of such process variants are contained, for example, in allprior art publications mentioned in relation to the dehydrogenationcatalysts.

A comparatively detailed description of dehydrogenation processessuitable according to the invention is also contained in CatalyticalStudies Division, Oxidative Dehydrogenation and AlternativeDehydrogenation Processes, Study Number 4192 OD, 1993, 430 FergusonDrive, Mountain View, Calif., 94043-5272 U.S.A.

Typical of partial heterogeneously catalyzed dehydrogenation of propaneis that it takes place endothermically, i.e. the heat (energy) necessaryfor establishing the required reaction temperature must be suppliedeither to the reaction gas beforehand and/or in the course of thecatalytic dehydrogenation.

Furthermore, owing to the high reaction temperatures required, it istypical of heterogeneously catalyzed dehydrogenations of propane thatsmall amounts of high-boiling high molecular weight organic compounds,including carbon, are formed and are deposited on the catalyst surfaceand thus deactivate the latter. In order to minimize thisdisadvantageous phenomenon, the propane to be passed over the catalystsurface for the catalytic dehydrogenation at elevated temperatures canbe diluted with steam. Under the resulting conditions, carbon depositedis partly or completely eliminated by the principle of the gasificationof coal.

Another possibility for eliminating deposited carbon compounds is topass an oxygen-containing gas through the dehydrogenation catalyst fromtime to time at elevated temperatures and thus more or less to burn offthe deposited carbon. Suppression of the formation of carbon deposits ishowever also possible by adding molecular hydrogen to the propane to bedehydrogenated catalytically, before it is passed at elevatedtemperatures over the dehydrogenation catalyst.

Of course, it is also possible to add steam and molecular hydrogen as amixture to the propane to be dehydrogenated catalytically. The additionof molecular hydrogen to the catalytic dehydrogenation of propane alsoreduces the undesired formation of allene and acetylene as byproducts.

A suitable reactor form for the propane dehydrogenation is the fixed-bedtubular reactor or tube-bundle reactor. This means that thedehydrogenation catalyst is present as a fixed bed in a reaction tube orin a bundle of reaction tubes. The reaction tubes are heated bycombustion of a gas, for example a hydrocarbon, such as methane, in thespace surrounding the reaction tubes. It is advantageous to use thisdirect form of catalyst tube heating only over the initial about 20 to30% of the fixed bed and to heat up the remaining bed length to therequired reaction temperature by the radiant heat liberated in thecourse of the combustion. In this way, an almost isothermal reaction isachievable. Suitable internal diameters of the reaction tubes are fromabout 10 to 15 cm. A typical dehydrogenation tube-bundle reactorcomprises from 300 to 1000 reaction tubes. The temperature in theinterior of the reaction tube is from 300 to 700° C., preferably from400 to 700° C. Advantageously, the reaction gas is preheated to thereaction temperature before being fed to the tubular reactor.Frequently, the product gas mixture leaves the reaction tube at atemperature of from 50 to 100° C. lower. In the abovementionedprocedure, the use of oxidic dehydrogenation catalysts based on chromiumoxide and/or alumina is expedient.

Frequently, no diluent gas is present but substantially pure propane isemployed as starting reaction gas. The dehydrogenation catalyst, too, isgenerally used undiluted.

On the industrial scale, about three tube-bundle reactors would beoperated in parallel. Two of these reactors would as a rule be in thedehydrogenation mode while the catalyst load is regenerated in one ofthe reactors.

The above procedure is used, for example, in the BASF Linde propanedehydrogenation process known in the literature.

Furthermore, it is used in the steam active reforming (STAR) processwhich was developed by Phillips Petroleum Co. (cf. for example U.S. Pat.No. 4,902,849, U.S. Pat. No. 4,996,387 and U.S. Pat. No. 5,389,342). Thedehydrogenation catalyst used in the STAR process is advantageouslypromoter-containing platinum on zinc (magnesium) spinel as a support(cf. for example U.S. Pat. No. 5,073,662). In contrast to the BASF Lindepropane dehydrogenation process, propane to be dehydrogenated in theSTAR process is diluted with steam. A molar ratio of steam to propane inthe range of from 4 to 6 is typical. The operating pressure isfrequently from 3 to 8 atom and the reaction temperature is expedientlychosen to be from 480 to 620° C. Typical catalyst loadings with thetotal reaction gas mixture are from 0.5 to 10 h⁻¹.

The propane dehydrogenation can also be designed in the form of a movingbed. For example, the moving catalyst bed can be housed in a radial flowreactor. The catalyst moves therein slowly from top to bottom while thereaction gas mixture flows radially. This procedure is used, forexample, in the UOP Oleflex dehydrogenation process. Since the reactorsin this process are operated quasi-adiabatically, it is expedient tooperate a plurality of reactors connected in series (typically up tofour). This makes it possible to avoid excessively large differences inthe temperatures of the reaction gas mixture at the reactor entrance andat the reactor exit (in the case of the adiabatic mode of operation, thereaction gas starting mixture acts as a heat-transfer medium on whoseheat content the reaction temperature is dependent) and nevertheless toachieve attractive total conversions.

When the catalyst bed has left the moving-bed reactor, it is regeneratedand then reused. For example, a spherical dehydrogenation catalyst whichsubstantially comprises platinum on spherical alumina supports can beused as a dehydrogenation catalyst for this process. In the UOP variant,hydrogen is added to the propane to be dehydrogenated, in order to avoidpremature catalyst aging. The operating pressure is typically from 2 to5 atm. The molar hydrogen-to-propane ratio is expediently from 0.1 to 1.The reaction temperatures are preferably from 550 to 650° C. and thetime for which the catalyst is in contact with the reaction gas mixtureis chosen to be from about 2 to 6 h⁻¹.

In the fixed-bed process described, the catalyst geometry may likewisebe spherical or cylindrical (hollow or solid).

As a further process variant for the propane dehydrogenation,Proceedings De Witt, Petrochem. Review, Houston, Tex., 1992 a, N1,describes the possibility of heterogeneously catalyzed propanedehydrogenation in a fluidized bed, in which the propane is not diluted.

Expediently, two fluidized beds are operated side by side, one of whichis as a rule present in the regeneration state. The active material usedis chromium oxide on alumina. The operating pressure is typically from 1to 1.5 atm and the dehydrogenation temperature is as a rule from 550 to600° C. The heat required for the dehydrogenation is introduced into thereaction system by preheating the dehydrogenation catalyst to thereaction temperature. The operating pressure is usually from 1 to 2 atmand the reaction temperature is typically from 550 to 600° C. The abovedehydrogenation procedure is also known in the literature as theSnamprogetti-Yarsintez process.

As an alternative to the procedures described above, the heterogeneouslycatalyzed propane dehydrogenation in the substantial absence of oxygencan also be realized according to a process developed by ABB LummusCrest (cf. Proceedings De Witt, Petrochem. Review, Houston, Tex., 1992,P1). The heterogeneously catalyzed propane dehydrogenation processes inthe substantial absence of oxygen which have been described to date havein common the fact that they are operated at propane conversions of >30mol % (as a rule ≦60 mol %) (based on a single reactor pass). It isadvantageous that it is sufficient to achieve a propane conversion offrom ≧5 to ≦30 or ≦25 mol %. This means that the propane dehydrogenationcan also be operated at propane conversions of from 10 to 20 mol % (theconversions are based on a single reactor pass). This is due, interalia, to the fact that the remaining amount of unconverted propane isdiluted in the downstream oxidation step (c) with molecular nitrogen,which reduces the propionaldehdye and/or propionic acid byproductformation.

For realizing the abovementioned propane conversions, it is advantageousto carry out the propane dehydrogenation at an operating pressure offrom 0.3 to 3 atm. It is also advantageous to dilute the propane to bedehydrogenated with steam. Thus, on the one hand, the heat capacity ofthe water makes it possible to compensate some of the effect of theendothermic nature of the dehydrogenation and, on the other hand, thedilution with steam reduces the partial pressure of starting materialsand products, which has an advantageous effect on the equilibriumposition of the dehydrogenation. Furthermore, as stated above, thepresence of steam has an advantageous effect on the time-on-stream ofthe dehydrogenation catalyst. If required, molecular hydrogen may alsobe added as a further component. The molar ratio of molecular hydrogento propane is as a rule <5. With a comparatively low propane conversion,the molar ratio of steam to propane can accordingly be from >0 to 30,expediently from 0.1 to 2, advantageously from 0.5 to 1. The fact thatonly a comparatively small amount of heat is consumed in a singlereactor pass of the reaction gas and comparatively low reactiontemperatures are sufficient for achieving the conversion in a singlereactor pass proves to be advantageous for a procedure having a lowpropane conversion.

It may therefore be expedient to carry out the propane dehydrogenation(quasi) adiabatically with a comparatively low propane conversion, i.e.the reaction gas starting mixture is heated as a rule to a temperatureof from 500 to 700° C. (for example by direct firing of the surroundingwall) or from 550 to 650° C. Usually, a single adiabatic pass through acatalyst bed is then sufficient to achieve the desired conversion, thereaction gas mixture being cooled by from about 30 to 200° C. (dependingon conversion). The presence of steam as a heat-transfer medium isadvantageous even from the point of view of an adiabatic procedure. Thelower reaction temperature permits longer times-on-stream of thecatalyst bed used.

In principle, the propane dehydrogenation with comparatively low propaneconversion, whether adiabatically or isothermally operated, can becarried out both in a fixed-bed reactor and in a moving-bed orfluidized-bed reactor.

It is noteworthy that a single shaft furnace reactor in the form of afixed-bed reactor through which the reaction gas mixture flows axiallyand/or radially is sufficient for realizing said dehydrogenation, inparticular in adiabatic operation.

In the simplest case, this is a single closed reaction volume, forexample a container, whose internal diameter is from 0.1 to 10 m,possibly also from 0.5 to 5 m, and in which the catalyst bed isinstalled on a support apparatus (for example a grille). The reactionvolume which is loaded with catalyst and is heat-insulated in adiabaticoperation is flowed through axially by the hot, propane-containingreaction gas. The catalyst geometry may be spherical, annular orstrand-like. Since in this case the reaction volume is to be realized bya very economical apparatus, all catalyst geometries which have aparticularly low pressure drop are preferable. These are in particularcatalyst geometries which lead to a large cavity volume or are built upin a structured manner, for example honeycombs. In order to realizeradial flow of the propane-containing reaction gas, the reactor mayconsist, for example, of two cylindrical grilles present in a casing andmounted concentrically one inside the other, and the catalyst bed may bearranged in their annular gap. In the adiabatic case, the metal casingin turn would be thermally insulated.

The catalysts disclosed in DE-A 199 37 107, especially all thosedisclosed by way of example, are particularly suitable as a catalystload for the propane dehydrogenation with comparatively low propaneconversion in a single pass.

After a relatively long operating time, the abovementioned catalysts canbe regenerated, for example, in a simple manner by passing air dilutedwith nitrogen over the catalyst bed at from 300 to 600° C., frequentlyfrom 400 to 500° C., initially in the first regeneration stages. Thecatalyst loading with regeneration gas may be, for example, from 50 to10000 h⁻¹ and the oxygen content of the regeneration gas may be from 0.5to 20% by volume.

In further downstream regeneration stages, air can be used asregeneration gas under otherwise identical regeneration conditions. Itis expedient in application technology to flush the catalyst with inertgas (for example N₂) before its regeneration.

Thereafter, it is generally advisable to effect regeneration with puremolecular hydrogen or with molecular hydrogen diluted with inert gas(the hydrogen content should be ≧1% by volume) under an otherwiseidentical range of conditions.

The propane dehydrogenation with comparatively low propane conversion(≦30 mol %) can be operated in all cases at the same catalyst loadings(relating both to the reaction gas as a whole and the propane containedtherein) as the variants with high propane conversion (>30 mol %). Thisloading with reaction gas may be, for example, from 100 to 10000 h⁻¹,frequently from 100 to 3000 h⁻¹, i.e. often from about 100 to 2000 h⁻¹.

The propane dehydrogenation with comparatively low propane conversioncan be realized in a particularly elegant manner in a tray reactor.

This contains, spatially in succession, more than one catalyst bedcatalyzing the dehydrogenation. The number of catalyst beds may be from1 to 20, expediently from 2 to 8, but also from 3 to 6. The catalystbeds are preferably arranged radially or axially one behind the other.In terms of application technology, it is expedient to use the fixedcatalyst bed type in such a tray reactor.

In the simplest case, the fixed catalyst beds in a shaft furnace reactorare arranged axially or in the annular gaps of cylindrical grillesinstalled concentrically one inside the other. However, it is alsopossible to arrange the annular gaps in segments and, after radialpassage through a segment, to pass the gas into the next segment aboveor below.

Expediently, the reaction gas mixture is subjected to intermediateheating in the tray reactor on its way from one catalyst bed to the nextcatalyst bed, for example by passing it over heat exchanger ribs heatedwith hot gases or by passing it through pipes heated with hot combustiongases.

If the tray reactor is otherwise operated adiabatically, it issufficient for the desired propane conversions (≦30 mol %), particularlywith the use of the catalysts described in DE-A 199 37 107, inparticular the exemplary embodiments, to preheat the reaction gasmixture to a temperature from 450 to 550° C. before passing it into thedehydrogenation reactor and to keep it within this temperature rangeinside the tray reactor. This means that the total propanedehydrogenation is thus to be realized at extremely low temperatures,which proves to be particularly advantageous for the time-on-stream ofthe fixed catalyst beds between two regenerations.

It is even more elegant to carry out the intermediate heating describedabove by a direct method (autothermal procedure). For this purpose, alimited amount of molecular oxygen is added to the reaction gas mixture,before it flows through the first catalyst bed and/or between thedownstream catalyst beds. Depending on the dehydrogenation catalystused, limited combustion of the hydrocarbons contained in the reactiongas mixture, any coke deposited on the catalyst surface or coke-likecompounds and/or hydrogen formed in the course of the propanedehydrogenation and/or added to the reaction gas mixture is thuseffected (it may also be expedient in terms of application technology tointroduce into the tray reactor catalyst beds which are loaded withcatalyst which specifically (selectively) catalyzes the combustion ofhydrogen (and/or of hydrocarbon) (suitable catalysts of this type are,for example, those of U.S. Pat. No. 4,788,371, U.S. Pat. No. 4,886,928,U.S. Pat. No. 5,430,209, U.S. Pat. No. 5,530,171, U.S. Pat. No.5,527,979 and U.S. Pat. No. 5,563,314; for example, such catalyst bedscan be housed in the tray reactor so as to alternate with the bedscontaining the dehydrogenation catalyst)). The heat of reaction evolvedthus permits virtually isothermal operation of the heterogeneouslycatalyzed propane dehydrogenation in a quasiautothermal manner. As thechosen residence time of the reaction gas in the catalyst bed increases,a propane dehydrogenation with decreasing or substantially constanttemperature is possible, which permits particularly long times-on-streambetween two regenerations.

As a rule, an oxygen feed as described above should be effected so thatthe oxygen content of the reaction gas mixture is from 0.5 to 10% byvolume, based on the amount of propane and propene contained therein.Suitable oxygen sources are both pure molecular oxygen and oxygendiluted with inert gas, for example CO, CO₂, N₂ or noble gases, inparticular air. The resulting combustion gases generally have anadditional diluting effect and thus promote the heterogeneouslycatalyzed propane dehydrogenation.

The isothermal nature of the heterogeneously catalyzed propanedehydrogenation can be further improved by mounting closed internals(for example annular ones), advantageously but not necessarilyevacuated, in the spaces between the catalyst beds in the tray reactorbefore they are introduced. Such internals may also be placed in therespective catalyst bed. These internals contain suitable solids orliquids which evaporate or melt above a specific temperature and thusconsume heat and condense and thereby liberate heat where thetemperature falls below this temperature.

One possibility of heating the reaction gas mixture to the requiredreaction temperature in the propane dehydrogenation is also to combust apart of the propane and/or H₂ contained therein by means of molecularoxygen (for example, over suitable combustion catalysts having aspecific action, for example by simply passing over and/or passingthrough) and to effect heating to the desired reaction temperature bymeans of the heat of combustion thus liberated. The resulting combustionproducts, such as CO₂ and H₂O, and any N₂ accompanying the molecularoxygen required for the combustion advantageously form inert diluentgases.

According to the invention, it is also possible for propane unconvertedand optionally propene after carrying out step (c) and separating offthe desired product (acrolein and/or acrylic acid) to be subjected topropane dehydrogenation, which can be carried out as described above,and for the product gas mixture obtained after the propanedehydrogenation to be subjected again to step (a).

Where a propane dehydrogenation is carried out, propane is a possiblediluent gas in step (c).

It is also possible to add, in particular when a propane dehydrogenationis carried out, to gas B supplied to step (c) still pure propane and/orpropene.

If step (c) is carried out as a conversion of propene to acrylic acid,then the exit gas from the working-up also contains oxidizable secondarycomponents, for example carbon monoxide, formic acid, formaldehyde,acetic acid and small amounts of acrylic acid in addition to thecomponents not converted in the oxidation, i.e. propane, nitrogen andresidual oxygen. In a particularly preferred embodiment, these secondarycomponents are catalytically oxidized before a propane dehydrogenationwith the residual oxygen and, if required, with additional molecularoxygen, in order to heat up the gas before the dehydrogenation. Thisoxidation could be carried out in a postcombustion catalyst, such as aPd catalyst on an alumina support, for example RO-20/13 or RO-20/25(both from BASF).

Preferred embodiments of the invention are shown in FIGS. 1 to 7described below, which illustrate the invention without restricting it.

FIGS. 1 to 5 show schematic diagrams for carrying out preferredprocesses, in which, for simplification, not all feed and dischargestreams are shown. FIG. 1 shows an absorption and desorption stage 1, anoxidation stage 2, which is in the form of an oxidation of propene toacrolein and/or acrylic acid, and a working-up stage 3. In FIG. 1,propane and propene, if required with residual amounts of nitrogen, areabsorbed into a suitable absorbent in stage 1 from a mixture whichcontains propane, propene, hydrogen and oxides of carbon (carbonmonoxide and carbon dioxide) and possibly nitrogen and furtherhydrocarbons, and are desorbed from said absorbent by stripping withair. In this way, hydrogen, the oxides of carbon, further hydrocarbonsand nitrogen are removed. The stream containing propene and possiblepropane is then fed to the oxidation stage 2, in which propene isoxidized to acrolein and/or acrylic acid. After the oxidation 2, theproduct obtained is fed to the working-up 3. There, the desired productsacrolein and/or acrylic acid are isolated. The remaining unconvertedpropene and propane and oxides of carbon and any residues of nitrogenand oxygen are once again fed to the absorption and desorption stage 1.

In the further figures, identical reference numerals denote the same asin FIG. 1.

In contrast to FIG. 1, in FIG. 2 a propane dehydrogenation 4 isinstalled upstream and can be carried out with or without a supply ofoxygen. The gas mixture obtained in the propane dehydrogenation andcontaining hydrogen, oxides of carbon and possibly residues of nitrogenand hydrocarbons in addition to the propane and the propene is fed tothe absorption and desorption stage 1.

In contrast to FIG. 1, in FIG. 3 a propane oxidation stage 22, in whichpropane is oxidized to acrolein and/or acrylic acid, is present insteadof propene oxidation 2.

In FIG. 4, a propane dehydrogenation stage 4 with or without a supply ofoxygen is carried out after the working-up stage 3 and the gas mixtureobtained in this stage is recycled to the absorption and desorptionstage 1.

FIG. 5 shows a further preferred embodiment of the process, in which apropane dehydrogenation 5 with an oxygen supply is installed downstreamof the absorption and desorption stage 1.

FIGS. 6 and 7 show further preferred processes. The process of FIG. 6follows the process diagram of FIG. 4. In FIG. 4, three reactors arepresent in the propane dehydrogenation stage 4, in the first of which acarbon monoxide postcombustion (CO-PC) takes place before the propanedehydrogenation (PDH), while in the two downstream reactors a hydrogenpostcombustion (H2-PC) takes place before the propane dehydrogenation(PDH). These postcombustions serve for supplying energy for the propanedehydrogenation. The number of reactors in the propane dehydrogenationis not limited to three reactors. In the propane dehydrogenation 4,propane is fed in via line (30). Air can be fed in via line (6). The gasmixture obtained after the propane dehydrogenation is fed via a heatexchanger W and a compressor V in stage 1 to an absorption column K1 anda desorption column K2. After the desorption in K2, the absorbent isrecycled to absorption column K1. Unabsorbed gases are removed from theprocess as waste gas (33), if necessary via an incineration plant E. Thestream containing separated off propane and/or propene is fed to theoxidation stage 2, which is shown here with four oxidation reactors.However, the number of oxidation reactors is not limited to this number.The desired product acrolein and/or acrylic acid is then worked up instage 3 and is taken off via line (31). Unconverted propane and/orpropene is recycled as recycle gas via line (32), together with theother gaseous components not separated off here in the absorption, tothe propane dehydrogenation 4.

In FIG. 7, recycle gas (1) from the working-up stage 3, which gas isobtained at from 10 to 90° C. and from 0.8 to 5 bar and can be furthercompressed to pressures of from 2 to 10 bar, for example with the aid ofa compressor V0, is heated, in a heat exchanger W1 countercurrently tothe reaction gas (2) from the propane dehydrogenation (PDH) 4, totemperatures of from 100 to 650° C. In the case of FIG. 7, the statedpressure in bar relates here and below to absolute pressure.

Suitable compressors are all suitable embodiments which are known to aperson skilled in the art and are mentioned in more detail below.

The recycle gas stream (1) contains from about 40 to 80% by volume ofN₂, from about 1 to 5% by volume of CO₂, from 0.1 to 2% by volume of CO,from 2 to 5% by volume of O-₂, from 0.5 to 25% by volume of H₂O, furtheroxidation byproducts and from about 5 to 40% by volume of unconvertedpropane and from about 0.1 to 3% by volume of unconverted propene.Before or after the heating-up, fresh propane (3) and preferably wateror steam (4) are mixed with the gas before it is passed into the PDH 4.Suitable fresh propane is any available propane-containing gas orliquid. However, propane sources such as industrial propane (>90%, inparticular >95%, particularly preferably >98%, propane content with asmall C₄ ⁺ fraction) or pure propane (>99% propane content) areadvantageous. The molar ratio of water or steam to propane in the gasstream is from 0.05 to 5, preferably from 0.1 to 2. It may beadvantageous to mix with this gas stream additionally H₂ (5), air (6) oran O₂-containing gas and further components capable of exothermicconversion in the PDH, for example CO or CO/H₂ mixtures, such assynthesis gas. The purity of these gases is not subject to anyrestriction. The exothermic nature of the oxidation of the combustiblecomponents in the PDH serves for covering the endothermic nature of thePDH reaction, so that less additional heat, and in the most advantageouscase no additional heat has to be supplied from outside for covering theenthalpy of reaction for the PDH.

The PDH is operated at from 0.3 to 10, preferably from 1 to 5, bar andfrom 350 to 700° C., preferably from 400 to 600° C. Possible reactorsfor the PDH are all embodiments known to a person skilled in the art,for example axial-flow apparatuses, such as tray reactors, and alsoapparatuses having a plurality of catalyst beds which are arranged inthe form of a hollow cylinder and having radial flow, or a plurality ofindividual apparatuses, for example column-type, cake-type or sphericalapparatuses. The number of reactors in the PDH is not limited to 3.Preferably, a plurality of individual apparatuses is used since theintermediate feeding of further gases is possible in a simple mannerthereby and moreover individual catalyst beds can be treated in aparticular manner, for example regenerated, separately from the othersduring operation. For this purpose, for example, the reactor containingthe catalyst bed to be regenerated is isolated from the main gas streamby suitable shut-off elements, for example slide valves, valves or flapswhich are present in the connecting lines between the reactors, and thegases required for regeneration, for example N₂, H₂, lean air or air orO₂-rich gases, are then passed over the catalyst bed and deposits areremoved from the catalyst. The remaining reactors, a total number ofwhich may be from 1 to 20, preferably from 2 to 5, are still fed withthe main gas stream and produce mainly the desired product propene.

In the PHD reactors, the catalyst layers may rest on grilles, beds ofinert material or similar support apparatuses known to a person skilledin the art. The form of the catalyst is not subject to any restrictions.Forms such as chips, spheres, extrudates, rings, cylinders or structuredpackings and monoliths may be used. Those geometries which provide asmall pressure drop are advantageous.

For the distribution of the gas fed in over the catalyst bed, gasdistributors known to a person skilled in the art, for example sievetrays, ring distributors or manifolds, and irregular beds or structuredpackings, for example static mixers, may be used.

A plurality of catalyst layers having different functions may bearranged in the PDH reactors. If a plurality of catalyst layers areused, it is advantageous to arrange, before the PDH catalyst layer, oneor more catalyst layers over which preferably, for example, H₂, COand/or a further oxidizable component which is not propene or propanecan be oxidized (CO—PC or H₂—PC). However, it is also possible todispense with additional catalyst layers upstream of the PDH catalystlayer if the PDH catalyst performs this function or propane lossesthrough oxidation with propane are economically acceptable.

The PDH catalysts can be operated with from 100 to 20000, preferablyfrom 500 to 10000, more preferably from 1000 to 10000, l(S.T.P.) ofpropane per liter of catalyst bed per hour. The gas space velocity forcatalysts which oxidize, for example, predominantly CO or H₂ and to alesser extent propane or propene is usually from 5000 to 30000 l(S.T.P.)of gas per liter of catalyst bed per hour.

The propane conversion in the PDH is from 10 to 60%, preferably from 20to 50%, at propane selectivities of from 80 to 99.5%, frequently from 88to 96%. The conversion of the feed gases CO, H₂ or other combustiongases is advantageously complete. The conversion of H₂ formed during thePDH is from 1 to 99%, often from 10 to 80%, frequently from 30 to 70%,depending on the propane conversion.

The reaction gas (2) from the PDH contains from about 20 to 60% byvolume of N₂, from about 1 to 5% by volume of CO₂, from 0.5 to 45% byvolume of H₂O, from about 5 to 40% by volume of propane, from about 1 to20% by volume of propene, from about 1 to 20% by volume of H₂ andfurther byproducts, for example ethane, ethene, methane and C₄ ⁺.

The reaction gas (2) from the PDH is obtained at from 400 to 650° C.,more advantageously from 450 to 600° C., and from 0.3 to 10, moreadvantageously from 1 to 5, bar. It is cooled, countercurrently to therecycle gas (1), to temperatures which are at least 5° C., better atleast 50° C., and preferably at least 100° C., above the inlettemperature of the recycle gas (1). The gas stream (3) is then furthercooled in one or more stages to about 10 to 60° C., depending on thetemperature on emergence from the countercurrent heat exchanger W1.

In the multistage cooling, the cooling in W2 can be effected by steamgeneration or by air cooling and the cooling in W3 by air, water orbrine cooling, depending on the temperature level. Depending on thepressure, temperature and H₂O content in the gas streams (3) to (5),water condenses and is separated from the gas stream (5) in theseparator A1. Suitable gas separators are all embodiments known to aperson skilled in the art and suitable for this purpose.

The cooled and if necessary partly dewatered gas stream (6) is thencompressed to pressures from the pressure on emergence from theseparator A1 to 50 bar. The compression can be effected either in onestage or in a plurality of stages with or without intermediate cooling.Suitable compressors V1 are all embodiments known to a person skilled inthe art and suitable for this purpose, for example reciprocating androtary compressors, screw-type compressors, diaphragm-type compressors,rotary multi-vane compressors, turbo compressors, centrifugalcompressors and rotary piston blowers and centrifugal blowers; however,turbo compressors or centrifugal compressors are preferably used.Criteria for choosing the compressors are both the pressure increase andthe amount of gas stream to be compressed. In the multistage compressionwith intermediate cooling, water and possibly further condensablecomponents condense during the intermediate cooling and can be separatedfrom the gas stream during or after the intermediate cooling asdescribed above, before the gas stream is fed to the next compressorstage. The gas stream (7) compressed to the final pressure can be cooledagain as described above in one or more stages, it being possible onceagain for water and any other condensable substances to be separatedfrom the gas streams (7) to (9).

The compressor V1 may be operated both by means of electric motors andby means of steam or gas turbines. The choice depends on theinfrastructure conditions. Frequently, driving by means of steam turbineproves most economical.

The sum of the condensed streams, for example (11)—after a pressureincrease—and (12), is recirculated to the PDH, to the extent that isrequired for covering the H₂O-to-propane ratio before entry into thePDH, and the remainder is discharged and if necessary incinerated. Thecondensate stream (13) can be vaporized before recirculation ordischarge or can be subjected to a further treatment, for example apurification, before recirculation.

The gas stream (10) is then fed to the absorption column K1, in whichpropane and/or propene are separated from the gas stream. Here, the gasstream (10) is brought into contact with an absorbent, which takes upthe C₃ fraction and may take up further components. Suitable absorbentsare all substances known to a person skilled in the art, the absorbentsdescribed above preferably being used. The gas stream (10) is preferablyfed countercurrently to the absorbent in a plurality of stages. Theabsorption can be effected at from 10 to 150° C., better at from 20 to80° C., preferably at from 30 to 60° C., and at from 1 to 50, better atfrom 3 to 30, preferably at from 5 to 20, bar.

Suitable absorbers K1 are all embodiments known to a person skilled inthe art, as described, for example, in Thermische Trennverfahren; KlausSattel, VCH, 1988 (ISBN 3-527-28636-5). Columns having internals arepreferable. Suitable internals are likewise all embodiments known to aperson skilled in the art, for example sieve trays, dual-flow trays,bubble trays, tunnel trays, lattice trays, valve trays or irregularbeds, for example comprising rings (for example from Raschig), Pallrings, Intalox saddles, Berl saddles, super saddles, toroidal saddles,Interpack packing or wire mesh rings and structured packings (forexample Sulzer-Kerapak or Sulzer packing BX or CY, or, for example,packings from Montz and packings from other manufacturers). Ralu-Pak250.YC from Raschig is particularly suitable. Internals which permithigh liquid loading or irrigation density, for example unstructured bedsor structured packings, are preferable. The possible irrigation densityshould be greater than 50, preferably greater than 80, m³ of liquid perm² of free cross-sectional area per hour. The internals may be eithermetallic, ceramic or of plastic or may consist of a compositioncomprising a plurality of materials. What is important in the choice ofthe material for the beds and packings is that the absorbent thoroughlywets these internals.

The ratio of the streams between the absorptive (24) fed to theabsorption and gas stream (10) follows the thermodynamic requirementsand depends on the number of theoretical plates, the temperature, thepressure, the absorption properties of the absorbent and the requireddegree of separation. Ratios of from 1:1 to 50:1, in particular from 1:1to 30:1, preferably from 1:1 to 20:1, in kg/kg, with from 1 to 40, inparticular from 2 to 30, preferably from 5 to 20, theoretical plates,are usual. The definition of a theoretical plate appears in thetechnical literature, for example “Thermische Trennverfahren”, KlausSattel, VCH, 1988, (ISBN 3-527-28636-5).

The gas stream (14) in which the concentration of propane and/or propenehas been reduced can be fed to a quench stage in order, if required, toreduce absorbent losses. The mode of operation of a quench is explainedin more detail below in the description of the desorption stage.

After leaving any quench stage, the gas stream (14) can be let down.Letting down can be effected either in one stage or in a plurality ofstages by throttling without energy recovery, or in one or more stagesin a gas turbine T1 with recovery of mechanical energy. In the case ofthe recovery of mechanical energy, it may be necessary to heat up thegas stream (14) before it is passed into the turbine. The gas stream canbe heated up both directly by catalytic and noncatalytic oxidation ofcombustible and oxidizing components contained in the gas stream or fedin from outside, and by indirect heat supply with the aid of steam orexternal firing. The mechanical energy obtained during let-down can beused directly as a concomitant or main means for driving one of thecompressors, preferably V1, or for generating electric power.

After the let-down, the waste gas stream (15) obtained can, depending onits purity, be fed to a catalytic or noncatalytic waste gas incinerationor discharged directly into the atmosphere.

The absorbent stream (16) laden predominantly with propane (from 2 to30% by volume) and/or propene (from 2 to 30% by volume) and possiblyfurther components (for example CO₂, C₂ ⁻, C₄ ⁺, H₂O), is let down ifnecessary and then fed to the desorption column K2. The let-down can beeffected both without recovery of the mechanical energy in one or morestages and with recovery of mechanical energy (for example in a turbineor a centrifugal pump operating in reverse). Moreover, it may be usefulto heat up the stream (16) prior to desorption. This heating-up ispreferably effected by means of countercurrent heat exchange with thestream (17) in W6. In addition, it may be useful to heat up stream (16)over and above this.

The desorption in K2 of propane and/or propene can be carried out bydistillation, by simple flash or by stripping. The desorption issupported by reducing the pressure to 0.1 to 10, in particular to 1 to5, preferably to 1.5 to 3, bar.

If the desorption is effected by distillation, the separation step canbe carried out on the basis of all knowledge known to a person skilledin the art. A particularly simple embodiment of the desorption is theone-stage flash or flash evaporation of the laden solvent in anapparatus suitable for this purpose. It may be expedient to heat thestream (16) to 20 to 300° C., in particular to 40 to 200° C., preferablyto 50 to 150° C., prior to flashing. The apparatus should be designed sothat both the thermodynamic separation between propane or propene andthe solvent and the fluid dynamic separation between gas and liquid takeplace readily. The apparatus may have, for example, a cylindrical orspherical shape, as well as other designs known to a person skilled inthe art. If the apparatus is of a cylindrical shape, the cylinder may beeither upright or horizontal. Viewed vertically, the feed to the flashapparatus is as a rule between the gas discharge and the liquiddischarge. In the simplest case, the apparatus has no additionalinternals. For better thermodynamic separation, the internals such asthose known to a person skilled in the art for distillation, absorptionand stripping can be installed in the apparatus, in particular thosedescribed in the text above for the absorption. For better fluid dynamicseparation, internals such as those known to a person skilled in the artfor gas/liquid separation, for example knitted fabrics, deflector platesor the like, may additionally be integrated in the flash apparatus.Moreover, the flash apparatus may contain apparatuses which permit theintroduction of heat, for example heated pipe coils or heated walls.

If, as in the present case, air or a similar stripping medium (forexample steam, N₂, fresh propane or a further gas required in theprocess) is available, it is expediently used for supporting the flashprocess.

A special embodiment for this purpose is the multistage stripping of thevolatile components propane and/or propene with the starting gas stream(25) for the oxidation (of course, as in the one-stage case, alladditionally absorbed components from the stream (10) and any substancesformed are also stripped according to their volatility). In the simplestcase, the starting gas stream (25) is the air required for oxidizing thepropene to acrolein or acrylic acid. The compression of the air or ofthe starting gas stream can be effected both before and after thedesorption. However, the starting gas stream (25) may also containrecycled gas from the acrylic acid process and steam, fresh propane orfurther blanketing gaseous component in addition to the air. It isparticularly advantageous if the starting gas stream (25) is fedcountercurrently or crosswise to the liquid absorbent during thedesorption. In the case of countercurrent flow, the desorption apparatusor desorber may be designed in the same way as the absorption columndescribed in the text above.

The cross-flow may be expedient if the explosion range is passed throughduring the desorption. This is the case when the starting gas stream(25) is a lean gas mixture with respect to the combustion tendency andthe gas stream (18) laden with propane and/or propene is a rich gasmixture with respect to the combustion inclination after the desorption.The gas mixture is defined as being lean in this context when thecontent of combustible substances is too low to be ignitable and a gasmixture is defined as rich in this context when the content ofcombustible substances is too high to be ignitable.

In the case of cross-flow, the total starting gas stream is notintroduced into the bottom but is divided into part-streams andintroduced at a plurality of suitable points along the desorptioncolumn, this being done in such a way that an ignitable gas mixture isnot present at any point in the desorption apparatus. The desorptioncolumn may be arranged vertically or horizontally.

A further possibility for overcoming the explosion problem in thedesorption of combustible components with O₂-containing gas streams, isto mix the starting gas stream, prior to entry into the desorptioncolumn, with a substance (for example propane, propene, methane, ethane,butane, H₂O, etc.) in such a way that the starting gas mixture is itselfrich prior to entry into the desorption column. However, it is alsopossible to split the starting gas stream and to pass a propane- orpropene-free starting gas into the bottom of the desorption column inorder to achieve very good depletion of propane and/or propene in theabsorptive (17) and to pass a starting gas which, for example, isenriched with propane and/or propene into that region of the desorptioncolumn in which an ignitable gas mixture can form.

After any countercurrent heat exchange (W6) and a pressure increase withthe aid of a pump (P1), the absorbent stream (17) depleted in propaneand/or propene can be further cooled in one or more stages (for examplein W7) and fed via line (24) back to the absorber K1.

In general, the multistage desorption may take place at all pressuresand temperatures.

However, pressures which are lower, and temperatures which are higher,than those in the absorption are advantageous. In the present case,pressures of from 1 to 5, in particular from 2 to 3, bar andtemperatures of from 20 to 200° C., in particular from 30 to 100° C.,particularly preferably from 35 to 70° C., are desirable.

The ratio of absorbent stream (17) to starting gas stream (25) followsthe thermodynamic requirements and depends on the number of theoreticalplates, the temperature, the pressure and the desorption properties ofthe absorbent and the required degree of separation. Ratios of from 1:1to 50:1, in particular from 5:1 to 40:1, preferably from 10:1 to 30:1,in kg/kg with from 1 to 20, in particular from 2 to 15, preferably from3 to 10, theoretical plates are usual.

In general, the starting gas stream laden with propane and/or propenecan be fed without further treatment to the oxidation stages 2. However,it may be expedient to feed the starting gas stream, prior to theoxidation, to a further process stage in order, for example, to reducethe losses of concomitantly stripped absorbent. The separation of theabsorbent from the laden starting gas stream for the oxidation can becarried out by all process steps known to a person skilled in the art.One possible embodiment is the quenching of the laden starting gasstream with water. In this case, the absorbent is washed out of theladen starting gas stream with water. This washing or quenching can becarried out at the top of the desorption column over a liquid collectingtray or in a separate apparatus. Internals such as those known to aperson skilled in the art for distillation, absorption and desorptionand as described in the text above for the absorption can be installedin the quench apparatus for supporting the separation effect. The sameapplies to the quench apparatus where it is designed as a separateapparatus.

After the water has washed the absorbent out of the starting gas streamladen with propane and/or propene, the water/absorbent mixture (19) canbe fed to a phase separation D1, and the treated starting gas stream(18), after possible preheating, can be fed to the propene oxidationstage 2.

The phase separation can be carried out in all embodiments known to aperson skilled in the art, as also used, for example, in liquid/liquidextraction. In the simplest case, these are horizontal or verticalelongated apparatuses with or without internals, in which the organicabsorbent phase separates from the quench water. The diameter-to-lengthratio here may be from 1:1 to 1:100, in particular from 1:1 to 1:10,preferably from 1:1 to 1:13. The apparatus may be flooded or may beoperated using a gas cushion. For better isolation of the organicabsorbent phase, the apparatus can be equipped with a dome from whichthe organic phase can be taken off. For supporting the phase separation,all internals known to a person skilled in the art for this purpose, forexample knitted fabrics, wound cartridges or deflector plates, may beinstalled. Of course, rotating phase separators, for examplecentrifuges, may also be used.

After the phase separation, the absorptive (20) separated off can berecycled to the desorption. The quench water can, if required, be cooledor heated in a heat exchanger (W9) before reentering the quenchapparatus. Advantageously, large amounts of water are circulated withthe aid of a pump (P2). Suitable irrigation densities in the quenchapparatus are greater than 30, in particular greater than 50, preferablygreater than 80, but less than 1 000, in particular 500, preferably lessthan 300, m³ of water per m² of free cross-sectional area of the quenchapparatus per hour.

The water losses during quenching can be covered by condensate (21) aswell as by dilute acid solution (22) from the acrylic acid preparationprocess. In order to avoid increasing concentrations, a part of thecirculation quench water can be removed as a purge stream (23) and fedto the incineration plant or to another treatment for disposal (forexample in a wastewater treatment plant).

If the gas stream (18) for the propene oxidation 2 has a temperature of<90° C., in particular <70° C. at pressures of from 1 to 3 bar, it maybe expedient also to add water to this stream. This can be effected byadmixing steam or by saturating the stream (18) in a water saturator ina manner known to a person skilled in the art. The gas stream treated inthis manner has a composition of from 30 to 70% by volume of N₂, from 5to 20% by volume of O₂, from 2 to 15% by volume of propene, from 2 to40% by volume of propane and from 0.5 to 25% by volume of H₂O andcontains further components, for example CO₂, methane, ethane, etheneand C₄ ⁺. It can be fed to the oxidation 2, which can be carried out asdescribed above or as disclosed in the patent literature. The propene oracrolein oxidation can be carried out in salt bath reactors, for examplefrom Deggendorfer-Werft according to the prior art or in other reactortypes. An air feed or steam feed may or may not once again take placebetween the first oxidation stage to acrolein and the second oxidationstage to acrylic acid. The higher C₃ content in the gas (18) with theoxidation 2 in any case requires removal of the heat of reaction fromthe reaction space. Propene loadings of from 5 to 350, in particularfrom 90 to 300, preferably from 100 to 250, l(S.T.P.) of propene perliter of catalyst bed per hour are suitable.

The separation of the acrylic acid from the reaction gas (26) of theoxidation 2 in stage 3 can be carried out as described above using, forexample, a high-boiling solvent, such as a mixture of diphyl anddimethyl phthalate, or by absorption in water and by fractionalcondensation.

The purification of the acrylic acid can be effected by stripping anddistillation or by azeotropic distillation or by crystallization.

The process described in FIG. 7 is suitable both for retrofitting of allexisting plants for the production of acrolein and/or acrylic acid andin combination with new acrylic acid plants.

Surprisingly, it was found that, in spite of the usually expectedresidues of absorbent in the gas B, no problems occurred with theoxidation or with the oxidation catalyst. Moreover, no problems wereobserved with any oxidation products which may form from the absorbentduring the oxidation. Where problems occur with residues of absorbent,which as a rule is not the case when hydrocarbons having a high boilingpoint, in particular paraffins, are used as absorbent, said absorbentcan be removed, for example by a water quench or by adsorption.

It was therefore surprising that absorption can be used in the novelprocess. In contrast to the adsorption used in the Japanese publicationJP-A-10 36311, the absorption used here for propane and/or propene issubstantially easier and more economical to handle.

Furthermore, the present invention has the advantage that existingplants for the preparation of acrolein and/or acrylic acid which usepropene as a starting material can be converted in an advantageousmanner to the more economical propane as a starting material.

The example which follows and which describes the preferred embodimentof the novel process illustrates the invention.

EXAMPLE

Acrylic acid is prepared by a process as shown in FIG. 7. The referencenumerals used below therefore relate to FIG. 7.

2090 l(S.T.P.)/h of the recycle gas (1) from the working-up stage 3,which is obtained at a temperature of 30° C. and a pressure of 1.2 bar,are compressed to 2.0 bar with the aid of a compressor V0 and heated to450° C. in a heat exchanger W1 countercurrently to the reaction gas (2)from the propane dehydrogenation (PDH). The stated pressure in barrelates here and below in this example to the absolute pressure.

The recycle gas stream (1) contains 60.3% by volume of N₂, 1.2% byvolume of CO₂, 0.5% by volume of CO, 3.4% by volume of O₂, 1.9% byvolume of H₂O, 32.2% by volume of propane and 0.4% by volume of propeneand further oxidation byproducts. Before the heating-up, 170 l(S.T.P.)/hof fresh propane (3) and steam (4) are mixed with the recycle gas streambefore it is passed into the PDH. The fresh propane used is industrialpropane (>98% propane content with 100 ppm by weight of C₄ ⁺fraction).The molar ratio of steam to propane in the gas stream is 0.5.

The gas mixture is fed to the 1st reactor of 4 reactors. The internaldiameter of the reactors is 50 mm. The reactors are designed in such away that they can be operated autothermally. Each reactor contains a 110mm high catalyst bed comprising extrudates (d=3 mm, 1=5 mm).

In the 4 reactors, propane undergoes 20% conversion at a propeneselectivity of 92%.

The reaction gas (2) from the PDH contains 44.9% by volume of N₂, 2.7%by volume of CO₂, 16.9% by volume of H₂O, 24.0% by volume of propane,5.8% by volume of propene, 5.5% by volume of H₂ and small amounts offurther byproducts, for example ethane, ethene, methane and C₄ ⁺.

The reaction gas (2) from the PDH is obtained at 520° C. and 1.5 bar andis cooled to 30° C. On average, about 350 g of water (11) per hourcondense.

The cooled gas stream (6) is then compressed in one stage in a pistoncompressor to 7.5 bar and is cooled again to 30° C. The condensed water(12) is combined with the condensation (11) and is partly vaporized andmixed with the recycle gas stream (1). A further part (stream (21)) isfed to the quench. The remainder is discharged.

2 340 l(S.T.P.)/h of the gas stream (10) are then passed into the bottomof the absorption column K1 (metal wall, internal diameter=80 mm, columnlength 3 m). 60% of the volume of the absorption column are filled withpacking elements from Montz (Montz-Pak type B1).

The gas stream (10) contains 53.9% by volume of nitrogen, 3.3% by volumeof carbon dioxide, 0.4% by volume of water, 28.8% by volume of propane,7.0% by volume of propene, 6.6% by volume of hydrogen and small amountsof further byproducts, for example ethane, ethene, methane and C₄₊.

35 kg/h of low-C₃ tetradecane (24) from the desorption column are passedat 30° C. to the top of the absorption column K1.

The waste gas stream (14) still contains 1150 ppm by volume of propaneand 750 ppm by volume of propene. The waste gas stream (14) from theabsorption column is let down to ambient temperature and ambientpressure, respectively, via a pressure control valve and thenincinerated.

The laden absorbent stream (16) is removed from the bottom of the columnK1, let down to 2.4 bar via a pressure control valve and fed to the topof the desorption column K2.

The desorption column K2 has the same dimensions as the absorptioncolumn K1 and is loaded in the same manner with packings.

1310 l(S.T.P.) of compressed air at 2.45 bar and 30° C. are passed intothe bottom of the desorption column. The desorption column isthermostated at 40° C.

The exit gas from the desorption column (2190 l(S.T.P.)/h) contains30.7% by volume of propane, 7.4% by volume of propene, 12.3% by volumeof O₂, 46.4% by volume of N₂, 1.5% by volume of H₂O and 1.6% by volumeof CO₂ and small residues of tetradecane and is passed into a quenchapparatus, which is located in the top of the desorption column K2.

The bottom discharge of the desorption column K2 is transported via thepump P1 and the heat exchanger W7 to the top of the absorption columnK1.

The quench apparatus is a metal column likewise having an internaldiameter of 80 mm and is equipped with internals of the same type asthose in the absorption column K1. The water quench is operated at 30°C. In the quench apparatus, about 120 l of water per hour are sprayedonto the bed. A two-phase liquid mixture is removed from the bottom ofthe quench apparatus and is passed into a phase separator. The phaseseparator is a horizontal container having a diameter of 200 mm and alength of 500 mm and contains a fine knitted wire fabric which isinstalled in the first third of the phase separator, in the direction offlow. The aqueous phase removed from the phase separator D1 is pumpedback to the top of the quench apparatus. On average, about 1 g oftetradecane per hour is removed from the phase separator and passed intothe tetradecane storage vessel. The water losses during quenching arecompensated by condensation water (21).

The exit gas stream from the water quench is heated to 200° C. before itis fed to the two-stage oxidation.

The oxidation takes place in model tubes having an internal diameter of26 mm and a length of 4 m. The first model tube is filled with 2.7 m ofa catalyst as described in EP-A-0 575 879 and the second model tube isfilled with 3 m of a catalyst as described in EP-A-0 017 000. 315l(S.T.P.) of fresh air are additionally passed per hour between thefirst and second oxidation stage.

The isolation of the acrylic acid from the reaction gas (26) of theoxidation and the purification of said acrylic acid are effected asdescribed in EP-A-0 982 289.

According to this process, on average 440 g of crude acrylic acid (27)comprising >99.5% of acrylic acid are obtained per hour.

1. A process for the preparation of at least one of acrolein and acrylicacid from at least one of propane and propene, the process comprisingthe following steps: (a) separating at least one of propane and propenefrom a gas mixture A containing at least one of propane and propene byabsorption in an absorbent, (b) separating at least one of propane andpropene from the absorbent to give a gas B containing at least one ofpropane and propene and (c) oxidizing the gas B obtained in stage (b) toform at least one of acrolein and acrylic acid, wherein noheterogeneously catalyzed dehydrogenation of propane without a supply ofoxygen is carried out between steps (b) and (c), and wherein theseparating in step (b) is carried out by stripping with at least one ofa pressure change and a temperature change, using at least one of steam,air and an oxygen/nitrogen mixture, and wherein during the separating(a) the gas mixture A is brought into contact with the absorbent at apressure of from 1 to 50 bar, and wherein the stripping (b) is carriedout at a pressure of from 0.1 to 10 bar.
 2. The process as claimed inclaim 1, wherein the gas mixture A further comprises at least onecomponent selected from the group consisting of hydrogen, nitrogen andoxides of carbon.
 3. The process as claimed in claim 1, wherein at leastone C₈-C₂₀-alkane or C₈-C₂₀-alkene is used as the absorbent in step (a).4. The process as claimed in claim 1, wherein, in step (c), propene isoxidized to at least one of acrolein and acrylic acid.
 5. The process asclaimed in claim 1, wherein, in step (c), propane is oxidized to atleast one of acrolein and acrylic acid.
 6. The process as claimed inclaim 5, wherein a multimetal oxide material of the formula (I)MoV_(b)M¹ _(c)M² _(d)O_(n)  (I) where M¹ is at least one of Te or Sb, M²is at least one element selected from the group consisting of Nb, Ta, W,Ti, Al, Zr, Cr, Mn, Ga, Fe, Ru, Co, Rh, Ni, Pd, Pt, La, Bi, B, Ce, Sn,Zn, Si and In, b is from .0.01 to 1, c is from >0 to 1, d is from >0 to1, and n is a number which is determined by the valency and frequency ofthe elements other than oxygen in (I), is used as the catalyst foroxidizing propane in step (c).
 7. The process as claimed in claim 1,wherein the gas mixture A used in step (a) has the composition of a gasmixture which is obtained by at least one of homogeneous orheterogeneously catalyzed dehydrogenation of propane to propane.
 8. Theprocess as claimed in claim 7, wherein the propane dehydrogenation iscarried out with a supply of oxygen.
 9. The process as claimed in claim1, wherein, after step (c) has been carried out, unconverted propane andoptionally propene is subjected to a propane dehydrogenation and theproduct mixture obtained is subjected to step (a) again.
 10. The processas claimed in claim 1, wherein step (c) is carried out directly afterstep (b).
 11. The process as claimed in claim 1, wherein after step (b)and before step (c) a water quench is carried out for separatingabsorbent.
 12. A process for the preparation of at least one of acroleinand acrylic acid from at least one of propane and propene, the processcomprising the following steps: (a) separating at least one of propaneand propene from a gas mixture A containing at least one of propane andpropene by absorption in an absorbent, (b) separating at least one ofpropane and propene from the absorbent to give a gas B containing atleast one of propane and propene and (c) oxidizing the gas B obtained instage (b) to form at least one of acrolein and acrylic acid, wherein noheterogeneously catalyzed dehydrogenation of propane without a supply ofoxygen is carried out between steps (b) and (c), wherein, after step (c)has been carried out, at least one of unconverted propane andunconverted propene is separated off according to steps (a) and (b) andis recycled to step (c), and wherein the separating in step (b) iscarried out by stripping with at least one of a pressure change and atemperature change, using at least one of steam, air and anoxygen/nitrogen mixture, and wherein during the separating (a) the gasmixture Ais brought into contact with the absorbent at a pressure offrom 1 to 50 bar, and wherein the stripping (b) is carried out at apressure of from 0.1 to 10 bar.
 13. The process as claimed in claim 12,wherein, in step (c), propene is oxidized to at least one of acroleinand acrylic acid.
 14. The process as claimed in claim 12, wherein, instep (c), propane is oxidized to at least one of acrolein and acrylicacid.
 15. The process as claimed in claim 12, wherein the gas mixture Aused in step (a) has the composition of a gas mixture which is obtainedby at least one of homogeneous or heterogeneously catalyzeddehydrogenation of propane to propene.
 16. The process as claimed inclaim 12, wherein, after step (c) has been carried out, unconvertedpropane and optionally propene is subjected to a propane dehydrogenationand the product mixture obtained is subjected to step (a) again.
 17. Theprocess as claimed in claim 12, wherein step (c) is carried out directlyafter step (b).
 18. The process as claimed in claim 12, wherein afterstep (b) and before step (c) a water quench is carried out forseparating absorbent.
 19. The process as claimed in claim 1, furthercomprising: dehydrogenating propane by heterogeneous catalysis in thepresence of oxygen.
 20. The process as claimed in claim 12, furthercomprising: dehydrogenating propane by heterogeneous catalysis in thepresence of oxygen.
 21. The process as claimed in claim 1, wherein theoxidizing is carried out in the presence of an oxidation catalyst withsubstantially no decrease in the activity of the oxidation catalyst. 22.The process as claimed in claim 1, wherein the absorbent is an organicsolvent having a boiling point of from 200 to 350° C.
 23. The process asclaimed in claim 1, wherein the absorbent is tetradecane.
 24. Theprocess as claimed in claim 1, further comprising: quenching gas B afterthe separating (b).
 25. The process as claimed in claim 24, whereinquenching includes spraying water into the gas B to form a two phasemixture and separating an aqueous phase of the two phase mixture from anorganic phase of the two phase mixture.
 26. The process as claimed inclaim 25, wherein the organic phase of the two phase liquid comprisesthe absorbent.
 27. The process as claimed in claim 12, wherein theoxidizing is carried out in the presence of an oxidation catalyst withsubstantially no decrease in the activity of the oxidation catalyst. 28.The process as claimed in claim 12, wherein the absorbent is an organicsolvent having a boiling point of from 200 to 350° C.
 29. The process asclaimed in claim 12, wherein the absorbent is tetradecane.
 30. Theprocess as claimed in claim 12, further comprising: quenching gas Bafter the separating (b).
 31. The process as claimed in claim 30,wherein the quenching includes spraying water into the gas B to form atwo phase mixture and separating an aqueous phase of the two phasemixture from an organic phase of the two phase mixture.
 32. The processas claimed in claim 31, wherein the organic phase of the two phaseliquid comprises the absorbent.
 33. The processes claimed in claim 1,wherein any hydrogen present in the gas mixture A is not present afterthe separating (b).
 34. The processes claimed in claim 12, wherein anyhydrogen present in the gas mixture A is not present after theseparating (b).
 35. The process as claimed in claim 1, wherein thestripping is carried out using air.
 36. The process as claimed in claim12, wherein the stripping is carried out using air.
 37. The process asclaimed in claim 1, wherein the stripping (b) is carried out at apressure of from 1 to 5 bar, and the gas mixture A is contacted with theabsorbent at a temperature of from 30 to 50° C.
 38. The process asclaimed in claim 12, wherein the stripping (b) is carried out at apressure of from 1 to 5 bar, and the gas mixture A is contacted with theabsorbent at a temperature of from 30 to 50° C.